Systems, methods, and devices for methane conversion via gas recycling

ABSTRACT

In a first stage of a methane conversion system, at least some methane (CH 4 ) in an input gas flow stream can be converted into C 2  hydrocarbons, hydrogen gas (H 2 ), and aromatics to provide a first processed stream. The conversion can be direct non-oxidative methane conversion (DNMC). At least some of the aromatics can be removed from the first processed stream to provide a second processed stream. In a second stage of the methane conversion system, at least some of the H 2  can be removed from the second processed stream to provide a recycle stream. The recycle stream can be returned to the first stage of the methane conversion system for further conversion of methane and removal of aromatics and H 2  products.

CROSS-REFERENCE TO RELATED APPLICATION(S)

The present application claims the benefit of U.S. ProvisionalApplication No. 63/189,672, filed May 17, 2021, entitled “System,Devices, and Methods for Gas Conversion,” which is incorporated byreference herein in its entirety.

STATEMENT REGARDING FEDERALLY SPONSORED RESEARCH

This invention was made with government support under CBET1264599awarded by the National Science Foundation (NSF). The government hascertain rights in the invention.

FIELD

The present disclosure relates generally to conversion of methane, andmore particularly, to conversion of methane into liquid aromatics, forexample, by direct non-oxidative methane conversion (DNMC).

BACKGROUND

Conversion of natural gas into liquid fuels can allow for easier andless expensive transport from remote extraction sites. In conventionalsystems, methane (CH₄), which is the main constituent of natural gas,can be converted to value-added hydrocarbons via the steam reforming ofCH₄ to produce syngas (CO+H₂), followed by the Fischer-Trøpsch process;however, such systems result in low carbon and energy efficiencies.Other systems have employed direct conversion routes, such as oxidativecoupling of methane (OCM) or direct non-oxidative methane conversion(DNMC). Compared with other approaches, DNMC can be more selective,given its unique capability in forming C₂+ hydrocarbons (C₂ hydrocarbonsand aromatics) and H₂ while circumventing intermediate energy intensivestep. In addition, DNMC can use methane as the only reactant, which canbe beneficial for a modular unit at a remote extraction site. Yet, inconventional DNMC systems, the conversion can be hindered bythermodynamic limitations, as well as the high rate of catalystdeactivation from coke formation.

Embodiments of the disclosed subject matter may address one or more ofthe above-noted problems and disadvantages, among other things.

SUMMARY

Embodiments of the disclosed subject matter system provide systems,methods, and devices for methane conversion via gas recycling. In someembodiments, a first stage can provide direct non-oxidative methaneconversion (DNMC) to convert at least some methane into C₂ hydrocarbons,hydrogen gas (H₂), and aromatics. An aromatics separation device,downstream of the first stage, can be used to separate (or remove) atleast some of the produced aromatics, for example, by condensing thearomatics. A second stage, downstream of the aromatics separationdevice, can then be used to separate (or remove) at least some of theproduced H₂ from the remaining methane, C₂ hydrocarbons, and aromatics,for example, by permeating hydrogen ions through a membrane. An outputof the second stage can then be recycled back to the input of the firststage for reprocessing, for example, to convert additional methane andremove additional aromatics and H₂.

In some embodiments, the second stage can operate at a differenttemperature than the first stage, for example, to avoid formation ofcoke in the second stage that could otherwise impede H₂ removal viamembrane permeation. In some embodiments, providing a stage for DNMCseparate from the stage for H₂ separation can circumvent thethermodynamic limitations encountered by conventional DNMC (e.g., byoperating at a higher temperature) and alleviate catalyst deactivation(e.g., by avoiding coke formation), while also producing high purity H₂and improving aromatics liquid yield. In some embodiments, the secondstage can be configured for autothermal operation, for example, by usingthe heat generated from a combustion reaction between separated H₂(e.g., permeated through a membrane to a sweep gas volume) and oxygen(e.g., oxygen gas (O₂) or an oxygen-containing compound) in a sweep gas.

Alternatively, in some embodiments, the first stage comprises anintegrated membrane reactor for performing DNMC and H₂ separation, andthe second stage comprises another integrated membrane reactor forperforming further DNMC and H₂ separation. For each of the integratedmembrane reactors, a sweep gas comprising O₂ or an oxygen-containingcompound can be flowed through a sweep gas volume on a permeate-side ofthe membrane opposite to the gas flow volume where methane is converted.Oxygen ions can back-diffuse from the permeate-side to theretentate-side of the membrane, where the oxygen reacts with carbonproduced by the DNMC to prevent, or at least reduce, coking on themembrane. Moreover, each of the integrated membrane reactors can beconfigured for autothermal operation by using the heat generated from acombustion reaction between permeated H₂ and oxygen in the sweep gas.

In a representative embodiment, a methane conversion system can comprisefirst and second membrane reactors, first and second gas supplies, anaromatics separation device, and a recycle line. The first membranereactor can comprise a first gas flow volume, a second gas flow volume,and a first membrane separating the first gas flow volume from thesecond gas flow volume. The first gas flow volume can have a firstcatalyst therein. The first gas supply can be coupled to the second gasflow volume and can be constructed to provide a first sweep gas to thesecond gas flow volume. The first sweep gas can comprise O₂ or anoxygen-containing compound. The aromatics separation device can beconnected to receive a first processed stream from the first gas flowvolume. The second membrane reactor can comprise a third gas flowvolume, a fourth gas flow volume, and a second membrane separating thethird gas flow volume from the fourth gas flow volume. The third gasflow volume can have a second catalyst therein. The second gas supplycan be coupled to the fourth gas flow volume and can be constructed toprovide a second sweep gas to the fourth gas flow volume. The secondsweep gas can comprise O₂ or an oxygen-containing compound. The recycleline can comprise one or more fluid conduits. The first gas flow volumeof the first membrane reactor can be connected to receive a recyclestream from the third gas flow volume of the second membrane reactor viathe recycle line.

The first reactor can be constructed to convert at least some CH₄ in aninput gas flow stream provided to the first gas flow volume of the firstreactor, so as to provide a first processed stream and such that aquantity of CH₄ in the first processed stream is less than that in theinput gas flow stream, the first processed stream comprising CH₄, C₂hydrocarbons, and aromatics. The C₂ hydrocarbons are acetylene (C₂H₂),ethylene (C₂H₄), ethane (C₂H₆), or any combination of the foregoing, andthe aromatics are benzene (C₆H₆), toluene (C₇H₈), naphthalene (C₁₀H₈),or any combination of the foregoing. The first membrane can beconstructed such that at least some H₂ is removed from the first gasflow volume by hydrogen ions permeating through the first membrane intothe second gas flow volume and such that oxygen ions permeate throughthe first membrane from the second gas flow volume into the first gasflow volume so as to reduce coking of the first membrane. The firstreactor can be constructed for autothermal operation via an exothermicreaction between the permeated hydrogen in the second gas flow volumeand the O₂ or oxygen-containing compound in the second gas flow volume.

The aromatics separation device can be constructed to remove at leastsome aromatics from the received first processed stream, so as toprovide a second processed stream comprising CH₄ and C₂ hydrocarbons,and to provide a first output stream comprising the removed at leastsome aromatics. A quantity of the aromatics in the second processedstream can be less than in the first processed stream. The secondreactor can be constructed to convert at least some CH₄ in the secondprocessed stream provided to the third gas flow volume of the secondreactor, so as to provide a recycle stream to the recycle line and suchthat a quantity of CH₄ in the recycle stream is less than in the secondprocessed stream. The second processed stream can comprise C₂hydrocarbons and aromatics. The second membrane can be constructed suchthat at least some H₂ is removed from the third gas flow volume byhydrogen ions permeating through the second membrane into the fourth gasflow volume and such that oxygen ions permeate through the secondmembrane from the fourth gas flow volume into the third gas flow volumeso as to reduce coking of the second membrane. The second reactor can beconstructed for autothermal operation via an exothermic reaction betweenthe permeated hydrogen in the fourth gas flow volume and the O₂ oroxygen-containing compound in the fourth gas flow volume.

In a representative embodiment, a methane conversion system can comprisea first reactor, an aromatics separation device, a second reactor, and arecycle line. The first reactor can have an inlet and an outlet. Thearomatics separation device can have an inlet, a first outlet, and asecond outlet. The inlet of the separation device can be connected toreceive a first processed stream from the outlet of the first reactor.The second reactor can have a first gas flow volume, a second gas flowvolume, and a membrane separating the first gas flow volume from thesecond gas flow volume. The first gas flow volume can be connected toreceive a second processed stream from the first outlet of the aromaticsseparation device. The recycle line can comprise one or more fluidconduits, and the inlet of the first reactor can be connected to receivea recycle stream from first gas flow volume via the recycle line.

The first reactor can be constructed to convert at least some CH₄ in aninput gas flow stream provided to the inlet of the first reactor, so asto provide to the outlet of the first reactor the first processed streamand such that a quantity of CH₄ in the first processed stream is lessthan that in the input gas flow stream. The first processed stream cancomprise CH₄, C₂ hydrocarbons, H₂, and aromatics. The aromaticsseparation device can be constructed to remove at least some aromaticsfrom the first processed stream provided to the inlet of the aromaticsseparation device, so as to provide to the first outlet of the aromaticsseparation device a second processed stream comprising CH₄, C₂hydrocarbons, and H₂, and to provide to the second outlet of thearomatics separation device a first output stream comprising the removedat least some aromatics. A quantity of the aromatics in the secondprocessed stream can be less than in the first processed stream. Thesecond reactor can be constructed to remove at least some H₂ from thesecond processed stream, which is provided to the first gas flow volume,into the second gas flow volume via the membrane, so as to provide tothe recycle line a recycle stream comprising CH₄ and C₂ hydrocarbons. Aquantity of the H₂ in the recycle stream can be less than that in thesecond processed stream.

In another representative embodiment, a method can comprise convertingat least some CH₄ in an input gas flow stream into C₂ hydrocarbons, H₂,and aromatics, thereby providing a first processed stream comprisingCH₄, C₂ hydrocarbons, H₂, and aromatics. A quantity of CH₄ in the firstprocessed stream can be less than that in the input gas flow stream. Themethod can further comprise removing at least some aromatics from thefirst processed stream, thereby providing a first output streamcomprising the removed at least some aromatics and a second processedstream comprising CH₄, C₂ hydrocarbons, and H₂. A quantity of thearomatics in the second processed stream can be less than that in thefirst processed stream. The method can also comprise removing at leastsome H₂ from the second processed stream, thereby providing a recyclestream comprising CH₄ and C₂ hydrocarbons. A quantity of the H₂ in therecycle stream can be less than that in the second processed stream. Themethod can further comprise providing the recycle stream as at leastpart of the input gas flow stream.

Any of the various innovations of this disclosure can be used incombination or separately. This summary is provided to introduce aselection of concepts in a simplified form that are further describedbelow in the detailed description. This summary is not intended toidentify key features or essential features of the claimed subjectmatter, nor is it intended to be used to limit the scope of the claimedsubject matter. The foregoing and other objects, features, andadvantages of the disclosed technology will become more apparent fromthe following detailed description, which proceeds with reference to theaccompanying figures.

BRIEF DESCRIPTION OF THE DRAWINGS

The patent or application file contains at least one drawing executed incolor. Copies of this patent or patent application publication withcolor drawing(s) will be provided by the Office upon request and paymentof the necessary fee.

Embodiments will hereinafter be described with reference to theaccompanying drawings, which have not necessarily been drawn to scale.Where applicable, some elements may be simplified or otherwise notillustrated in order to assist in the illustration and description ofunderlying features. Throughout the figures, like reference numeralsdenote like elements.

FIGS. 1A-1B are simplified schematic diagrams of exemplary methaneconversion systems with gas recycling, according to one or moreembodiments of the disclosed subject matter.

FIG. 1C is a simplified schematic diagram of another methane conversionsystem with multiple aromatic separation stages, according to one ormore embodiments of the disclosed subject matter.

FIG. 1D is a simplified schematic diagram of another methane conversionsystem with multiple reactor stages, according to one or moreembodiments of the disclosed subject matter.

FIG. 2A is a simplified schematic diagram of another methane conversionsystem, according to one or more embodiments of the disclosed subjectmatter.

FIG. 2B is a simplified cross-sectional view of an exemplary fixed-bedreactor setup that can be used for the direct non-oxidative methaneconversion (DNMC) reactor of FIG. 2A, according to one or moreembodiments of the disclosed subject matter.

FIG. 2C is a simplified cross-sectional view of an exemplary tubularreactor setup that can be used for the hydrogen separator of FIG. 2A,according to one or more embodiments of the disclosed subject matter.

FIG. 2D is a simplified cross-sectional view of another exemplarytubular reactor setup that can be used for the hydrogen separator ofFIG. 2A, according to one or more embodiments of the disclosed subjectmatter.

FIG. 2E is a simplified cross-sectional view of an exemplary tubularreactor setup for autothermal operation that can be used for the DNMCreactor of FIG. 2A, according to one or more embodiments of thedisclosed subject matter.

FIG. 3 is a process flow diagram for a method of methane conversion withgas recycle, according to one or more embodiments of the disclosedsubject matter.

FIG. 4 is a simplified schematic diagram illustrating another exemplaryconfiguration for a gas-recycle system with DNMC reactor and hydrogenseparator, according to one or more embodiments of the disclosed subjectmatter.

FIG. 5 depicts a generalized example of a computing environment in whichthe disclosed technologies may be implemented.

FIGS. 6A-6B are graphs of H₂ permeation flux and H₂ removal efficiency,respectively, through a single SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3-δ)membrane tube as a function of test temperature and H₂ partial pressure.

FIGS. 7A-7B are graphs of H₂ permeation flux and H₂ removal efficiency,respectively, of the SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3-δ) membrane tube atdifferent temperatures as a function of time on stream.

FIGS. 8A-8D are graphs of feed composition and flow rates to the methanereactor at different cycle numbers for H₂ removal efficiencies by thehydrogen separator of 10% H₂ removal, 40% H₂ removal, 70% H₂ removal,and 100% H₂ removal, respectively.

FIGS. 9A-9D are graphs of methane conversion and product selectivityafter the methane reactor at different cycle numbers for H₂ removalefficiencies by the hydrogen separator of 10% H₂ removal, 40% H₂removal, 70% H₂ removal, and 100% H₂ removal, respectively.

FIGS. 10A-10D are graphs of overall methane conversion and productselectivity at different cycle numbers for H₂ removal efficiencies bythe hydrogen separator of 10% H₂ removal, 40% H₂ removal, 70% H₂removal, and 100% H₂ removal, respectively.

FIG. 11 is a graph of overall aromatic liquid yield from DNMC atdifferent cycle numbers for H₂ removal efficiencies by the hydrogenseparator of 10% H₂ removal, 40% H₂ removal, 70% H₂ removal, and 100% H₂removal, respectively.

FIG. 12A is a graph of methane conversion and product yield for afixed-bed reactor with Fe@SiO₂ catalyst and for a membrane reactor withdifferent sweep gases.

FIG. 12B is a graph of product selectivity for a fixed-bed reactor withFe@SiO₂ catalyst and for a membrane reactor with different sweep gases.

FIG. 13A is a graph of methane conversion and product yield from DNMC bya heat-exchange H₂-permeable membrane reactor as a function oftemperature.

FIG. 13B is a graph of product selectivity from DNMC by theheat-exchange H₂-permeable membrane reactor as a function oftemperature.

FIG. 13C is a graph of H₂ production and permeation, and H₂ removalpercentage in the DNMC by the heat-exchange H₂-permeable membranereactor as a function of temperature.

FIG. 13D is a graph comparing heat requirement for DNMC and heatreleased from combustion of H₂ permeate for different temperatures.

FIGS. 14A, 14D, and 14G are graphs of methane conversion and productyield as a function of O₂ sweep flow rate during DNMC by theheat-exchange H₂-permeable membrane reactor at 1303 K, 1273 K, and 1253K, respectively.

FIGS. 14B, 14E, and 14H are graphs of product selectivity as a functionof O₂ sweep flow rate during DNMC by the heat-exchange H₂-permeablemembrane reactor at 1303 K, 1273 K, and 1253 K, respectively.

FIGS. 14C, 14F, and 14I are graphs of heat requirement for DNMC and heatrelease from combustion of H₂ permeate as a function of O₂ sweep flowrate during DNMC by the heat-exchange H₂-permeable membrane reactor at1303 K, 1273 K, and 1253 K, respectively.

FIG. 15A is a graph of methane conversion and product yield for DNMC bythe heat-exchange H₂-permeable membrane reactor at 1303 K as a functionof time on stream.

FIG. 15B is a graph of heat requirement for DNMC and heat release fromcombustion of H₂ permeate during DNMC by the heat-exchange H₂-permeablemembrane reactor at 1303 K as a function of time on stream.

FIG. 16 is a graph illustrating a two-dimensional temperature profile ofthe heat-exchange H₂-permeable membrane reactor during operation thereofat 1273 K.

DETAILED DESCRIPTION General Considerations

For purposes of this description, certain aspects, advantages, and novelfeatures of the embodiments of this disclosure are described herein. Thedisclosed methods and systems should not be construed as being limitingin any way. Instead, the present disclosure is directed toward all noveland nonobvious features and aspects of the various disclosedembodiments, alone and in various combinations and sub-combinations withone another. The methods and systems are not limited to any specificaspect or feature or combination thereof, nor do the disclosedembodiments require that any one or more specific advantages be present,or problems be solved. The technologies from any embodiment or examplecan be combined with the technologies described in any one or more ofthe other embodiments or examples. In view of the many possibleembodiments to which the principles of the disclosed technology may beapplied, it should be recognized that the illustrated embodiments areexemplary only and should not be taken as limiting the scope of thedisclosed technology.

Although the operations of some of the disclosed methods are describedin a particular, sequential order for convenient presentation, it shouldbe understood that this manner of description encompasses rearrangement,unless a particular ordering is required by specific language set forthbelow. For example, operations described sequentially may in some casesbe rearranged or performed concurrently. Moreover, for the sake ofsimplicity, the attached figures may not show the various ways in whichthe disclosed methods can be used in conjunction with other methods.Additionally, the description sometimes uses terms like “provide” or“achieve” to describe the disclosed methods. These terms are high-levelabstractions of the actual operations that are performed. The actualoperations that correspond to these terms may vary depending on theparticular implementation and are readily discernible by one skilled inthe art.

The disclosure of numerical ranges should be understood as referring toeach discrete point within the range, inclusive of endpoints, unlessotherwise noted. Unless otherwise indicated, all numbers expressingquantities of components, molecular weights, percentages, temperatures,times, and so forth, as used in the specification or claims are to beunderstood as being modified by the term “about.” Accordingly, unlessotherwise implicitly or explicitly indicated, or unless the context isproperly understood by a person skilled in the art to have a moredefinitive construction, the numerical parameters set forth areapproximations that may depend on the desired properties sought and/orlimits of detection under standard test conditions/methods, as known tothose skilled in the art. When directly and explicitly distinguishingembodiments from discussed prior art, the embodiment numbers are notapproximates unless the word “about” is recited. Whenever“substantially,” “approximately,” “about,” or similar language isexplicitly used in combination with a specific value, variations up toand including 10% of that value are intended, unless explicitly statedotherwise.

Directions and other relative references may be used to facilitatediscussion of the drawings and principles herein but are not intended tobe limiting. For example, certain terms may be used such as “inner,”“outer,”, “upper,” “lower,” “top,” “bottom,” “interior,” “exterior,”“left,” right,” “front,” “back,” “rear,” and the like. Such terms areused, where applicable, to provide some clarity of description whendealing with relative relationships, particularly with respect to theillustrated embodiments. Such terms are not, however, intended to implyabsolute relationships, positions, and/or orientations. For example,with respect to an object, an “upper” part can become a “lower” partsimply by turning the object over. Nevertheless, it is still the samepart, and the object remains the same.

As used herein, “comprising” means “including,” and the singular forms“a” or “an” or “the” include plural references unless the contextclearly dictates otherwise. The term “or” refers to a single element ofstated alternative elements or a combination of two or more elementsunless the context clearly indicates otherwise.

Although there are alternatives for various components, parameters,operating conditions, etc. set forth herein, that does not mean thatthose alternatives are necessarily equivalent and/or perform equallywell. Nor does it mean that the alternatives are listed in a preferredorder, unless stated otherwise. Unless stated otherwise, any of thegroups defined below can be substituted or unsubstituted.

Unless explained otherwise, all technical and scientific terms usedherein have the same meaning as commonly understood to one skilled inthe art to which this disclosure belongs. Although methods and materialssimilar or equivalent to those described herein can be used in thepractice or testing of the present disclosure, suitable methods andmaterials are described below. The materials, methods, and examples areillustrative only and not intended to be limiting. Features of thepresently disclosed subject matter will be apparent from the followingdetailed description and the appended claims.

Overview of Terms

The following is provided to facilitate the description of variousaspects of the disclosed subject matter and to guide those skilled inthe art in the practice of the disclosed subject matter.

C₂ hydrocarbons: Compounds formed by the conversion of methane (CH₄) andincluding two carbons (C₂) and at least two hydrogens (H₂). In someembodiments, the C₂ hydrocarbons can include acetylene (C₂H₂), ethylene(C₂H₄), ethane (C₂H₆), or combinations thereof. In some embodiments, thepredominant C₂ hydrocarbon formed by the conversion of methane isethylene.

Aromatics: Hydrocarbons formed by the conversion of methane andlarger/heavier than the C₂ hydrocarbons. In some embodiments, thearomatics can include benzene (C₆H₆), toluene (C₇H₈), naphthalene(C₁₀H₈), or combinations thereof.

Fe@SiO₂: A catalyst formed of iron (Fe) and silica (SiO₂), where @denotes confinement, and characterized by lattice-confined single ironsites embedded within a silica matrix.

Autothermal operation: Operation of the system where the heat used todrive one reaction (e.g., a direct non-oxidative methane conversion(DNMC)) or operation (e.g., hydrogen permeation) is provided by asimultaneous (or substantially simultaneous) exothermic reaction. Insome embodiments, the exothermic reaction comprises combustion ofhydrogen and oxygen to form water, where the hydrogen is provided bypermeation of hydrogen ions through a membrane.

Oxygen-containing compound: A compound having at least one atom ofoxygen and capable of releasing the oxygen to react with hydrogen (e.g.,to form water). In some embodiments, the oxygen-containing compound cancomprise CO₂, H₂O, alcohols (e.g., methanol, ethanol, isopropanol,etc.).

Introduction

Direct non-oxidative methane conversion (DNMC) can be used to convertmethane (CH₄) (e.g., in natural gas) into liquid aromatics, which may beeasier to transport than CH₄ gas. In conventional systems, DNMC isconstrained by low CH₄ conversion and thus low aromatics product yielddue to thermodynamic limitations. In one or more embodiments of thedisclosed subject matter, a gas-recycle system is provided with a firststage (e.g., a methane conversion reactor) and a separate second stage(e.g., a hydrogen (H₂) membrane separator) to achieve high CH₄conversion (e.g., at least 40% after multiple cycles) and high aromaticsyield (e.g., at least 50% after multiple cycles and/or with a productselectivity toward aromatics of at least 90%) by circumventing thethermodynamic limitations.

For example, in some embodiments, the methane conversion stage (e.g.,first stage) can perform DNMC over a catalyst (e.g., Fe@SiO₂) in apacked reactor to produce C₂ products and aromatics, along with reactionproduct H₂. In some embodiments, the produced aromatics can then beremoved, for example, by passing the product stream through a condenser.In some embodiments, the produced H₂ can also be removed, for example,by passing through an H₂-permeable membrane separator (e.g., a secondstage). The unreacted CH₄, C₂ hydrocarbons, residual aromatics (if any),and residual H₂ (if any) in the product stream can then be sent back tothe methane conversion stage via a recycle loop, for example, for thenext round of DNMC reaction, aromatics removal, and H₂ removal.

In some embodiments, system performance can be further improved byemploying an autothermal reactor configuration for the first stage, thesecond stage, or both. When provided as part of the first stage, theDNMC reaction can be performed over a catalyst (e.g., Fe@SiO₂) in amembrane reactor to produce C₂ hydrocarbons, aromatics, and H₂ gas in aproducts volume (e.g., a first gas flow volume). The membrane reactorcan have an H₂-permeable membrane that separates the products volumefrom a sweep gas volume (e.g., a second gas flow volume), and a sweepgas can be flowed through the sweep gas volume. In some embodiments, thesweep gas can comprise oxygen gas (O₂) (e.g., air or a mixture of O₂ andwith one or more other gases, such as He gas) or an oxygen-containingcompound. Within the sweep gas volume, the permeated H₂ can react withthe oxygen in the sweep gas to produce heat for the endothermic DNMCwithin the products volume (e.g., at least some of the heat required forthe endothermic DNMC reaction, and preferably all of the heat required).In some embodiments, the membrane can allow for back diffusion of O₂(e.g., via permeation of oxygen ions) from the sweep gas volume to theproducts volume, thereby oxidizing carbon species therein into carbonmonoxide (CO) and thus reducing carbon deposition in the membranereactor.

When provided as part of the second stage, the H₂ removal can beperformed in a membrane reactor with an H₂-permeable membrane thatseparate a processed flow volume (e.g., a first gas flow volume) from asweep gas volume (e.g., a second gas flow volume). A sweep gas can beflowed through the sweep gas volume. In some embodiments, the sweep gascan comprise oxygen gas (O₂) (e.g., air or a mixture of O₂ and with oneor more other gases, such as He gas) or an oxygen-containing compound.Within the sweep gas volume, the permeated H₂ can react with the oxygenin the sweep gas to produce heat that maintains or raises a temperatureof the membrane reactor (e.g., to support H₂ permeation through themembrane).

In an exemplary embodiment, each of the first and second stages employsan integrated membrane reactor with respective catalyst, and by using asweep gas comprising O₂ or an oxygen-containing compound for eachreactor. For each membrane reactor, the DNMC reaction can be performedover a catalyst (e.g., Fe@SiO₂) in a membrane reactor to produce C₂hydrocarbons, aromatics, and H₂ gas in a products volume (e.g., a firstgas flow volume). The membrane reactor can have an H₂-permeable membranethat separates the products volume from a sweep gas volume (e.g., asecond gas flow volume). Within the sweep gas volume, the permeated H₂can react with the oxygen in the sweep gas to produce heat for theendothermic DNMC within the products volume (e.g., at least some of theheat required for the endothermic DNMC reaction, and preferably all ofthe heat required). In some embodiments, the membrane can allow for backdiffusion of O₂ (e.g., via permeation of oxygen ions) from the sweep gasvolume to the products volume, thereby oxidizing carbon species thereininto carbon monoxide (CO) and thus reducing carbon deposition in themembrane reactor. This can allow the reactors of the first and secondstages to operate at a higher temperature (e.g., the same temperaturefor both stages) for greater hydrogen permeation and higher aromaticyield, while avoiding, or at least reducing, membrane coking.

Exemplary Methane Conversion System Configurations

FIG. 1A shows an exemplary gas recycle system 101 for methaneconversion. In the illustrated example, the gas recycle system 101includes a first stage 111, a second stage 131, an aromatics separationdevice 114 arranged inline between the first and second stages, and arecycle line 136 for returning a processed stream back to the firststage 111 for further processing. The first stage 111 can include afirst integrated membrane reactor 107 configured for methane conversion(e.g., via DNMC) and hydrogen separation (e.g., by removing H₂ from aprocessed stream), and the second stage 130 can include a secondintegrated membrane reactor 127 configured for methane conversion andhydrogen separation. The recycling of a processed stream for furtherprocessing by the system can further improve process efficiency (e.g.,by increasing the amount CH₄ converted and/or the amount of aromaticsproduced).

In the illustrated example, an initial methane feed 102 is provided viafeed coupler 138 as an input gas flow stream 104 to the first membranereactor 107 for conversion. For example, the first membrane reactor 107can convert at least some methane in the input gas flow stream 104 to C₂hydrocarbons, hydrogen gas (H₂), and aromatics. In addition, the firstmembrane reactor 107 can remove at least some of the produced H₂ bytransport across (e.g., permeation through) a membrane. The resultingfirst processed stream 113 can thus include at least C₂ hydrocarbons andaromatics, as well as any unreacted methane and potentially hydrogen notremoved through the membrane, where the quantity of methane in the firstprocessed stream 113 is less than the initial quantity of methane in theinput gas flow stream 104.

In some embodiments, the first membrane reactor 107 can be anH₂-permeable membrane reactor, for example, according to the reactorconfigurations or constructions described in U.S. Pat. No. 10,525,407,issued Jan. 7, 2020, and entitled “Systems, Methods, and Devices forDirect Conversion of Methane,” which is incorporated by reference hereinin its entirety. For example, the first membrane reactor 107 cancomprise and/or define a first flow volume and a second flow volume,where the permeable membrane separates the first flow volume from thesecond flow volume. The first flow volume can receive the input gas flowstream 104, and the first processed stream 113 can be directed from anoutlet (or outlet end) of the first flow volume.

A catalyst can be provided in the first flow volume of the firstmembrane reactor 107. In one or more embodiments, the catalyst cancomprise metal elements doped (i.e., lattice doping) in the lattice ofamorphous-molten-state materials made from Si bonded with one or two ofelemental C, N or O, for example, SiO₂. In lattice doping, the dopantmetal elements exchange with the lattice elements in the doped materialssuch that the metal dopant elements are confined in the lattice of thedoped materials. For example, the amount of dopant metal can be between0.001 wt % and 10 wt % of the total weight of the catalyst. For example,the dopant metal elements can be one or more of Li, Na, K, Mg, Al, Ca,Sr, Ba, Y, La, Ti, Zr, Ce, Cr, Mo, W, Re, Fe, Co, Ni, Cu, Zn, Ge, In,Sn, Pb, Bi, Mn, such as Fe. For example, the catalyst can compriseFe@SiO₂, which has lattice-confined single iron sites embedded in thesilica matrix. The Fe@SiO₂ catalyst disclosed herein can be formedaccording to the fabrication method described in U.S. Pat. No.10,525,407, incorporated by reference above. Although the discussionabove is directed to the Fe@SiO₂ catalyst, embodiments of the disclosedsubject matter are not limited thereto. Rather, according to one or morecontemplated embodiments, other catalysts may be used, such as but notlimited to molybdenum/Zeolite Socony Mobil-5 (Mo/ZSM-5) or a noblemetal.

The first reactor 107 can be provided in a first heating module 108,which heats and/or maintains the reactor 107 at a first temperature. Thefirst heating module 108 can comprise a heater, a furnace, or both;however, alternative heating methodologies and configurations are alsopossible according to one or more contemplated embodiments. For example,instead of or in supplement to a heater of the first heating module 108,heat can be generated via an exothermic reaction between gasconstituents in the first reactor 107, such as between a sweep gas andpermeated hydrogen. In particular, a sweep gas comprising O₂ or anoxygen-containing compound can be provided on a permeate-side of themembrane. The permeated hydrogen can react with the oxygen to generateheat for the methane conversion in the first reactor 107.

Moreover, oxygen ions can back-diffuse through the membrane (e.g., fromthe permeate-side into the product volume) to oxidize carbon speciesproduced by the methane conversion (e.g., forming CO), thereby avoiding,or at least reducing, coking of the membrane. The resulting firstprocessed stream 113 may thus further include CO. By avoiding coking,the first reactor 107 can operate at a higher temperature, therebyoffering greater hydrogen permeation and higher aromatic yields (e.g.,≥1100 K, as suggested by FIGS. 13A-13B).

The first processed stream 113 can be directed to an aromaticsseparation device 114, which can be configured to remove at least someof the aromatics from the first processed stream 112. For example, thearomatics separation device 114 can comprise a condenser that condensesthe aromatics in the first processed stream 113 while retaining theremaining components (e.g., CH₄, C₂ hydrocarbons, CO, and H₂) in gaseousstate to form a second processed stream 125. In some embodiments, thearomatics separation device 114 can include other components forpre-processing, such as a heat exchanger, chiller, etc., and/or othercomponents for post-processing of removed aromatics, such as adistillation system, etc. The aromatics separation device 114 can beother than a condenser and/or employ separation techniques other thanselective condensation according to one or more contemplatedembodiments. In the illustrated example, the liquid aromatics can thenbe collected as a first output stream 118 for storage, transport, or useby module 120. For example, module 120 can comprise a storage containeror a conduit for conveying liquid aromatics to a storage or forsubsequent use. Alternatively, in some embodiments, module 120 can beomitted in favor of on-site use of the liquid aromatics.

The second processed stream 125 from the aromatics separation device 114can include at least C₂ hydrocarbons, and as well as any unreactedmethane and potentially hydrogen not removed by the first reactor 107and/or aromatics not removed by the separation device 114, where thequantity of aromatics in the second processed stream 125 is less thanthe initial quantity of aromatics in the first processed stream 112. Inthe illustrated example, the second processed stream 125 is directed tothe second reactor 127 for further methane conversion. For example, thesecond membrane reactor 127 can convert at least some methane in thesecond processed stream 125 to C₂ hydrocarbons, H₂, and aromatics. Inaddition, the second membrane reactor 127 can remove at least some ofthe produced H₂ by transport across (e.g., permeation through) amembrane. The resulting recycle stream 133 can thus include at least C₂hydrocarbons and aromatics, as well as any unreacted methane andpotentially hydrogen not removed through the membrane and/or CO, wherethe quantity of methane in the recycle stream 133 is less than theinitial quantity of methane in the second processed stream 125.

In some embodiments, the second membrane reactor 127 can be anH₂-permeable membrane reactor, for example, according to the reactorconfigurations or constructions described in U.S. Pat. No. 10,525,407,incorporated by reference above. For example, the second membranereactor 127 can comprise and/or define a third flow volume and a fourthflow volume, where the permeable membrane separates the third flowvolume from the fourth flow volume. The third flow volume can receivethe second processed stream 125, and the recycle stream 133 can bedirected from an outlet (or outlet end) of the third flow volume. Acatalyst can be provided in the third flow volume of the second membranereactor 127. For example, the catalyst can be the same or different thanthe catalyst of the first membrane reactor 107.

The second reactor 127 can be provided in a second heating module 128,which heats and/or maintains the reactor 127 at a second temperature.The second heating module 128 can comprise a heater, a furnace, or both;however, alternative heating methodologies and configurations are alsopossible according to one or more contemplated embodiments. For example,instead of or in supplement to a heater of the second heating module128, heat can be generated via an exothermic reaction between gasconstituents in the second reactor 127, such as between a sweep gas andpermeated hydrogen. In particular, a sweep gas comprising O₂ or anoxygen-containing compound can be provided on a permeate-side of themembrane. The permeated hydrogen can react with the oxygen to generateheat for the further methane conversion in the second reactor 127.

Moreover, oxygen ions can back-diffuse through the membrane (e.g., fromthe permeate-side into the product volume) to oxidize carbon speciesproduced by the methane conversion (e.g., forming CO), thereby avoiding,or at least reducing, coking of the membrane. The resulting recyclestream 133 may thus further include CO. By avoiding coking, the secondreactor 127 can operate at a higher temperature (e.g., a sametemperature as the first reactor 107), thereby offering greater hydrogenpermeation and higher aromatic yields (e.g., ≥1100 K, as suggested byFIGS. 13A-13B).

As noted above, the sweep gas for the first reactor 107 or the secondreactor 127 can comprise O₂ (e.g., air, O₂ gas alone, or O₂ gas combinedwith one or more other gases) and/or an oxygen-containing compound(e.g., CO₂, H₂O, and/or alcohol), for example, to combust with permeatedH₂ to provide heat supporting the methane conversion in the respectivereactor. In the illustrated example, the sweep gas is provided to thesecond flow volume of the first reactor 107 via a sweep gas inlet feed121, and a second outlet stream 133 can be directed from an outlet (oroutlet end) of the second flow volume of the first reactor. In addition,the sweep gas is provided to the fourth flow volume of the secondreactor 127 via a sweep gas inlet feed 122, and a third outlet stream134 can be directed from an outlet (or outlet end of the fourth flowvolume of the second reactor. The second outlet stream 133 and the thirdoutlet stream 134 can each contain water, for example, resulting fromthe combustion between oxygen in the sweep gas and the permeated H₂. Insome embodiments, the second and third outlet streams 133, 134 can besubjected to further processing (e.g., isolation of the H₂ and/or waterfrom the sweep gas), storage, and/or use.

In the illustrated example, the recycle stream 133 can be returned tothe inlet of the system 100 via recycle line 136, for re-processing in anext cycle, which can further convert additional methane that wasunreacted in the previous cycle. For example, the recycle line 136 canconnect between an outlet of the second reactor 127 and a feed coupler138. In some embodiments, the feed coupler 138 can combine the recyclestream 133 from the recycle line 136 with fresh methane from feed 102for processing as input gas flow stream 104. Alternatively oradditionally, in some embodiments, the feed coupler 138 can selectbetween the recycle stream 133 from the recycle line 136 and the methanefeed 102 for use as the input gas flow stream 104. For example, therecycle stream 133 can be used as the input gas flow stream 104 forrepeated cycles (e.g., 3-10 passes through the first stage 111,aromatics separation device 114, and the second stage 131) until amajority (e.g., — 50%), most (e.g., —70%), or substantially all (e.g.,—90% or more) of an initial methane batch has been converted, afterwhich the feed coupler 138 can be switched to provide a new batch ofmethane from methane feed 102.

System 100 can further include a controller 140 operatively coupled toone, some, or all of the illustrated components and configured tocontrol operation thereof. For example, the controller 140 can modifyflow rates, feed gas composition, sweep gas composition, and/ortemperature to regulate methane conversion efficiency and/or productselectivity. Gas flow lines within system 100 can include respective gasflow control and sensing module, which may include, for example, valves,temperature sensors, temperature controllers, pumps, mass flowcontrollers, and/or other devices to monitor and/or control thevariables of gas flow rates, reaction temperatures, and/or feed andsweep gas compositions to optimize or otherwise control methaneconversion product formation, as described herein. In some embodiments,the controller 140 can also control operation of components notillustrated, such as pumps, valves, switches, etc., to effect flow offluid (e.g., liquid or gas mixtures) through the system.

Alternatively, in some embodiments, methane conversion and hydrogenseparation can be provided in separate stages rather than performedwithin the same reactor. For example, FIG. 1B shows an exemplary gasrecycle system 100 for methane conversion. In the illustrated example,the gas recycle system 100 includes a first stage 110, a second stage130, an aromatics separation device 114 arranged inline between thefirst and second stages, and a recycle line 136 for returning aprocessed stream back to the first stage 110 for further processing. Insome embodiments, the first stage 110 can include a first reactor 106configured for methane conversion (e.g., via DNMC), and the second stage130 can include a second reactor 126 configured for hydrogen separation(e.g., by removing H₂ from a processed stream). For example, the firstreactor 106 can be a fixed-bed reactor (e.g., with Fe@SiO₂ catalyst,Mo/ZSM-5 catalyst, noble metal catalyst, or any other catalyst) and thesecond reactor 126 can be a H₂-permeable membrane separator.

The first reactor 106 can convert at least some methane in the input gasflow stream 104 to C₂ hydrocarbons, hydrogen gas (H₂), and aromatics.The resulting first processed stream 112 can thus include at least C₂hydrocarbons, H₂, and aromatics, as well as any unreacted methane, wherethe quantity of methane in the first processed stream 112 is less thanthe initial quantity of methane in the input gas flow stream 104. Thefirst processed stream 112 can be directed to an aromatics separationdevice 114, which can be configured to remove at least some of thearomatics from the first processed stream 112. For example, thearomatics separation device 114 can comprise a condenser that condensesthe aromatics in the first processed stream 112 while retaining theremaining components (e.g., CH₄, C₂ hydrocarbons, and H₂) in gaseousstate to form a second processed stream 124.

The second processed stream 124 from the aromatics separation device 114can include at least C₂ hydrocarbons and H₂, and as well as anyunreacted methane and potentially aromatics not removed by theseparation device 114, where the quantity of aromatics in the secondprocessed stream 124 is less than the initial quantity of aromatics inthe first processed stream 112. In the illustrated example of FIG. 1B,the second processed stream 124 is directed to the second reactor 126for hydrogen separation (also referred to herein as hydrogen removal orhydrogen isolation). The resulting product stream 132 can thus includeat least C₂ hydrocarbons, as well as unreacted methane and potentiallyaromatics not removed by the separation device 114 and H₂ not removed bythe second reactor 126, where the quantity of H₂ in the product stream132 is less than the initial quantity of H₂ in the second processedstream 124. Similar to FIG. 1A, the product stream 132 in FIG. 1B can bereturned to the inlet of the system 100 via recycle line 136 forre-processing in a next cycle, which can further convert additionalmethane that was unreacted in the previous cycle.

By decoupling the methane conversion reaction from the hydrogen removal,the processes can be separately optimized, for example, to operate atrespective temperatures that increase process efficiency (e.g.,enhancing CH₄ conversion and/or aromatics yield) and/or maintain systemoperability (e.g., by avoiding coke formation). Depending on thecatalyst employed in the first reactor 106, the first reactor 106 canoperate at a different temperature (e.g., higher when using Fe@SiO₂ orlower when using Mo/ZSM-5) than the second reactor 126. For example, thefirst temperature can be at least 1000 K (e.g., ≥1200 K) for Fe@SiO₂catalyst, and the second temperature can be less than or equal to 1100 K(e.g., ≤800 K).

Although only two stages are illustrated in FIGS. 1A-1B, embodiments ofthe disclosed subject matter are not limited thereto. Rather, in someembodiments, one or more intermediate reaction stages (with or withoutassociated aromatics separation devices) can be provided between aninitial methane conversion stage (e.g., first stage 110 or 111) and afinal methane conversion stage (e.g., second stage 131) or a finalH₂-separation stage (e.g., second stage 130) to provide furtherprocessing (e.g., additional conversion of methane and/or removal ofaromatics prior to hydrogen removal and recycling). For example, FIG. 1Cshows another gas recycle system 150 for methane conversion, where anintermediate stage 160 for further processing is provided inline afterthe aromatics separation device 114 and before second stage 131. Thesecond processed stream 125 from the aromatics separation device 114 isprovided to the intermediate reactor 156 of the intermediate stage 160.

In some embodiments, the intermediate reactor 156 can function similarto the first reactor 107, for example, to provide an additional stage ofmethane conversion and hydrogen separation. Alternatively, in someembodiments, the intermediate reactor 156 can function similar to thesecond reactor 127, for example, to provide an additional stage ofmethane conversion and hydrogen separation or similar to second reactor126, for example, to provide an additional stage of hydrogen separationwithout associated methane conversion. The intermediate reactor 156 canbe provided in a third heating module 158, which heats and/or maintainsthe reactor 156 at a third temperature. Depending on system operation,the third temperature may be the same as the first temperature (e.g.,when providing additional DNMC) or different than the first temperature(e.g., to provide improved hydrogen separation at higher temperatures).For example, the intermediate reactor 156 can have a membrane thatseparates a product flow volume from a sweep gas flow volume. The secondprocessed stream 124 can be provided to the product flow volume of theintermediate reactor 156 and can exit therefrom as intermediateprocessed stream 162. A sweep gas inlet feed 152 (e.g., similar to sweepgas inlet feed 122) can be provided to the sweep gas flow volume of theintermediate reactor 156 and can exit therefrom as outlet stream 154(e.g., similar to outlet stream 134).

In the illustrated example, the intermediate processed stream 162 isdirected to a second aromatics separation device 164 (e.g., similar toaromatics separation device 114), which can be configured to remove atleast some of the aromatics from the intermediate processed stream 162.For example, the aromatics separation device 164 can comprise acondenser that condenses the aromatics in the intermediate processedstream 162 while retaining the remaining components (e.g., CH₄, C₂hydrocarbons, and H₂) in gaseous state to form further processed stream174, which can then be directed to the second reactor 126. In theillustrated example, the liquid aromatics can then be collected as asecond output stream 168 (e.g., similar to output stream 118) forstorage, transport, or use by module 170 (e.g., similar to module 120).Although shown separately in FIG. 1C, in some embodiments, modules 120,170 can be combined, for example, as a single volume within a container,flow volumes connecting a common conduit, separate volumes within acommon structure, or any other configuration.

Although FIGS. 1A-1C illustrate only a single reactor in each stage,embodiments of the disclosed subject matter are not limited thereto.Rather, in some embodiments, a stage can include multiple reactors. Forexample, FIG. 1D shows another gas recycle system 180 for methaneconversion, where a first stage 182 (e.g., for methane conversion viaDNMC) includes multiple first reactors 184, 186 and a second stage 190(e.g., for methane conversion via DNMC or for H₂ separation) includesmultiple second reactors 194, 196. The first processed stream 113 outputfrom the ultimate first reactor 186 of the first stage 182 can bedirected to aromatics separation device 114, and the second processedstream 125 from the aromatics separation device 114 can be directed tothe initial second reactor 194 of the second stage 190.

In some embodiments, the first reactors 184, 186 can have similarstructures and/or operate similarly to each other, for example, eachsimilar to reactor 107 of FIG. 1A or reactor 106 of FIG. 1B.Alternatively or additionally, the first reactors 184, 186 can havedifferent structures (e.g., one fixed-bed and the other a permeablemembrane reactor). Each first reactor 184, 186 can be provided in arespective heating module 188 a, 188 b (e.g., similar to heating module108), which heat and/or maintain the respective reactor at a respectivetemperature. In some embodiments, the first reactors 184, 186 canoperate at a same temperature (e.g., at least 1000 K). Alternatively, insome embodiments, the reactors 184, 186 can operate at differenttemperatures. Although shown separately in FIG. 1D, in some embodiments,heating modules 188 a, 188 b can be combined together in a singlemodule, for example, when first reactors 184, 186 operate at the sametemperature. Moreover, although only two reactors are illustrated forfirst stage 182, any number of reactors can be provided in series and/orin parallel.

In some embodiments, the second reactors 194, 196 can have similarstructures and/or operate similarly to each other, for example, eachsimilar to reactor 127 of FIG. 1A or reactor 126 of FIG. 1B.Alternatively or additionally, the second reactors 194, 196 can havedifferent structures, for example, one as a permeable membrane separator(e.g., no catalyst) and the other as a permeable membrane reactor (e.g.,with catalyst). Similar to reactor 127 in FIG. 1A, reactor 194 can beprovided with a sweep gas inlet feed, and reactor 196 can be providedwith a sweep gas inlet feed 192 b. In some embodiments, a composition ofthe sweep gas provided to each reactor 194, 196 can be the same.Alternatively, in some embodiments, the sweep gas composition of theinlet feed for reactor 194 can be different from the sweep gascomposition of the inlet feed for reactor 196, for example, depending onthe operation and/or configuration of the corresponding second reactor.

Each second reactor 194, 196 can be provided in a respective heatingmodule 198 a, 198 b (e.g., similar to heating module 128), which heatand/or maintain the respective reactor at a respective temperature. Insome embodiments, the second reactors 194, 196 can operate at a sametemperature (e.g., less than 1000 K). Alternatively, in someembodiments, the reactors 194, 196 can operate at differenttemperatures. Although shown separately in FIG. 1D, in some embodiments,heating modules 198 a, 198 b can be combined together in a singlemodule, for example, when second reactors 194, 196 operate at the sametemperature. Moreover, although only two reactors are illustrated forstage 190, any number of reactors can be provided in series and/or inparallel.

Exemplary Reactor Configurations

FIG. 2A, shows a schematic of a gas recycle system 200 with a recycleloop. Methane from an initial methane feed 202 can be provided as aninput gas flow stream 204 to DNMC reactor 206 via feed coupler 226. TheDNMC reactor 206 can include a catalyst for performing DNMC at hightemperature, thereby producing a first processed flow stream 208comprising equilibrated gas mixtures of CH₄, H₂, C₂ products, andaromatics in a single step. In some embodiments, the DNMC reactor 206can be a fixed-bed reactor or an H₂-permeable membrane reactor. Forexample, FIG. 2B illustrates a fixed-bed reactor setup 230 that can beused for the DNMC reactor 206 of FIG. 2A. The fixed-bed reactor setup230 can include a reactor housing 232 (e.g., a tube, such as a quartztube) defining an internal flow volume 234 and a catalyst 236 (e.g.,Fe@SiO₂) disposed within the volume 234.

In another example, FIG. 2E illustrates a tubular reactor setup 290 thatcan be used for the DNMC reactor 206 of FIG. 2A. The tubular reactor 290can have a first gas volume 280 formed by the interior volume of aporous support tube 274. An H₂ permeable membrane (not separatelyillustrated) can be provided on a surface of the support tube 274, forexample, on the radially outer surfaces of the support tube 274. Themembrane can be constructed to allow protonic/electronic transportbetween the first and second gas volumes, such that H₂ produced fromprior methane conversion can be removed from the first gas volume 280 tothe second gas volume 282. Note that the H₂ transport through themembrane is via bulk diffusion, i.e., ion transport without applicationof an external electric field. For example, support tube can be aperovskite-type oxide having a formula of M′Ce_(1-z)Zr_(z)O_(3-δ), whereM′ is Sr or Ba, and z is between 0.01 and 0.3, inclusive, and themembrane can be a perovskite-type oxide having a formula ofM′Ce_(1-x-y)Zr_(x)M″_(y)O_(3-δ), where M′ is a least one of Sr and Ba;M″ is at least one of Ti, V, Cr, Mn, Fe, Co Ni, Cu, Nb, Mo, W, Pr, Nd,Pm, Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm and Yb; x is between 0.01 and 0.2,inclusive; and y is between 0.01 and 0.3, inclusive. In someembodiments, the porous support tube is SrCe_(0.8)Zr_(0.2)O_(3-δ), andthe H₂-permeable is a SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) film on thesupport tube. Other suitable materials for the permeable membrane and/orporous support can be found in, for example, U.S. Pat. No. 10,525,407,incorporated by reference above.

Feed gas can be provided to the first gas volume 280 via an inlet tube272 disposed within the porous support tube 274. Inlet tube 272 canconvey the input gas flow stream 204 to the first gas volume 280 andinto contact with catalyst 296 (e.g., Fe@SiO₂). As illustrated in FIG.2E, the porous support tube 274 can be provided with an end cap portion,which may be integral with the support tube 274 or a separate pieceadhered to the support tube 274. The porous support tube 274 is thusclosed at one end and opened at the other. The feed gas is thusredirected by the end cap back toward the inlet along the first gasvolume 280 and through the catalyst 296 to thereby form first processedstream 208. Alternatively, the porous support tube 274 can be open atboth ends, in which case the outlet can be disposed at end of thereactor 290 opposite from the inlet (e.g. similar to FIG. 2C).

The reactor 290 can have a second gas volume 282 formed by the annularspace between the porous support tube 274 and an outer enclosure 276,e.g., a quartz tube. Sweep gas 292 can be provided to the second gasvolume 282 and exits therefrom as outlet stream 294. In someembodiments, the sweep gas 292 can comprise an inert gas (e.g., N₂ orHe). Alternatively, in some embodiments, the sweep gas 292 can compriseOz gas or an oxygen containing compound. For example, the CH₄ can beconverted to C₂₊ products via the DNMC reaction in the reactor 290following the equation CH₄→3/52 C₆H₆+5/104 C₁₀H₈+7/104 C₂H₄+2/104C₂H₂+19/13 H₂ (ΔH>0). Meanwhile, outside of the tube 274, the O₂ in thesweep gas can react with H₂ permeate (e.g., H₂+½O₂→H₂O, ΔH<0) to produceheat for the endothermic DNMC within the reactor tube. The energybalance between endothermic DNMC and exothermic H₂ combustion reactionson opposite sides of the membrane (e.g., supported on tube 274) can beachieved, thereby providing autothermal operation of DNMC. In someembodiments, back diffusion of O₂ from the sweep side (e.g., volume 282)to the DNMC side (e.g., volume 280) can oxidize carbon species into CO,thereby alleviating carbon deposition in the reactor 290 and avoidingcoke formation.

Returning to FIG. 2A, the first processed stream 208 is directed tocondenser 210, where the liquid aromatics in the stream 208 are removedfrom the effluent gases, for example, to the first output stream 212.The remaining effluent gases (e.g., CH₄, H₂, and C₂ products) are thendirected as second processed stream 214 to an H₂ separator 216, whichcan have an H₂ permeable membrane. H₂ in the second processed stream 214can be transported across the membrane (e.g., via permeation of hydrogenions) in the H₂ separator 216 to produce pure H₂ via the membraneseparation. For example, FIG. 2C illustrates a reactor setup 240 thatcan be used for H₂ separator 216 of FIG. 2A. Reactor 240 can beconfigured as a tubular reactor, with an inner tube 246 defining a firstgas volume 244 for the second processed stream 214 and an annular space248 between the inner tube 246 and an outer tube 242 serving as a secondgas volume for sweep gas 218. A wall of the inner tube 246 can be formedwith the permeable membrane thereon, such that the reactor can beconsidered a flow-through tubular membrane reactor. Alternatively, anend-capped reactor setup 250 can be used for H₂ separator 216 of FIG.2A, for example, as shown in FIG. 2D. Reactor 250 can also be configuredas a tubular reactor with a first gas volume 252 formed by the interiorvolume of a porous support tube 254. An H₂ permeable membrane 256 can beprovided on a surface of the support tube 254, for example, on theradially outer surfaces of the support tube 254. In the examples ofFIGS. 2C-2D, the support tube, the membrane, or both can be formedsimilar to that described above for the reactor setup 290 of FIG. 2E.

Returning to FIG. 2D, second processed stream 214 can be provided to thefirst gas volume 252 via an inlet tube 260 disposed within the poroussupport tube 254. Inlet tube 260 can convey the second processed stream214 to the first gas volume 252. The porous support tube 254 can beprovided with an end cap portion, such that the tube 254 is closed atone end and opened at the other. The second processed stream is thusredirected by the end cap back toward the inlet along the first gasvolume 252, whereby H₂ can permeate through the membrane 256 to thesecond gas volume 258, to thereby form effluent stream 220 exiting viafirst outlet 262. The second gas volume 258 can be formed by the annularspace between the porous support tube 254 and an outer enclosure 268,e.g., a quartz tube. Sweep gas 218 can be provided to the second gasvolume 258 via inlet 264 and exits as output stream 222 via secondoutlet 266.

In some embodiments of the reactor 240 of FIG. 2C or the reactor 250 ofFIG. 2D, the sweep gas provided to the second gas volume can comprise aninert gas (e.g., such as nitrogen, helium, neon, argon, krypton, xenon,radon, or combinations thereof). Alternatively, in some embodiments, thesweep gas can comprise O₂ gas (e.g., air, O₂ gas alone, or O₂ gascombined with one or more other gases) or an oxygen containing compound(e.g., CO₂, H₂O, and/or alcohol). The oxygen in the sweep gas can reactwith hydrogen that has permeated through the membrane to the second gasvolume to produce heat within the reactor tube, for example, in a mannersimilar to that described above for reactor 290 of FIG. 2E.

Returning to FIG. 2A, the effluent 220 from the H₂ separator 216 canthen be recycled back to input gas flow stream 204, via recycle line224, for introduction into the DNMC reactor 206 in a subsequent cycle.By increasing the number of cycles, CH₄ conversion can be increasedsignificantly beyond the first pass equilibrium, and high aromaticliquid yields can be achieved. Moreover, the decoupling of CH₄activation and H₂ separation into two operation units can improve carbonconversion efficiency, since heavier aromatic hydrocarbon products couldhave grown under DNMC reaction conditions to coke, which would otherwisedeteriorate membrane performance and reduce CH₄-to-aromatics conversionefficiency.

Although FIGS. 2B-2E illustrate specific configurations for componentsof system 200, other reactor configurations and variations (e.g., toprovide multiple reactors operating in series, in parallel, or both, forexample, to scale up to process practical quantities of methane) arealso possible according to one or more embodiments of the disclosedsubject matter. For example, any of the membrane reactor configurationsdisclosed in U.S. Pat. No. 10,525,407, incorporated by reference above,can be used in place of the DNMC reactor 206, the H₂ separator 216,and/or any other reactor disclosed herein.

Methane Conversion Method

FIG. 3 is a process flow diagram for an exemplary method 300 of methaneconversion with gas recycle. The method 300 can initiate at processblock 302, where at least some methane in an input gas flow stream isconverted into C₂ hydrocarbons, H₂, and aromatics via interaction with acatalyst at an elevated temperature. For example, the methane conversioncan comprise DNMC. In some embodiments, the methane conversion can beperformed by or within one or more first reactors, such as any ofreactors 106, 184, 186, 206, 230, 290, and 410. The methane conversioncan occur at a first temperature, for example, at least 1000 K (e.g.,≥1200 K). In some embodiments, the conversion of process block 302 canproduce a first processed stream comprising the C₂ hydrocarbons, H₂,aromatics, and any unreacted methane.

The method 300 can proceed to decision block 304, where it is determinedif hydrogen should be removed prior to aromatics separation. If it isdesired to remove hydrogen, the method 300 can proceed from decisionblock 304 to process block 306, where a first sweep gas can be flowed ona first side (e.g., within a second gas flow volume) of a first membrane(e.g., of the first reactor) opposite to the input gas flow streamand/or the first processed stream (e.g., within a first gas flowvolume). In some embodiments, the first sweep gas can comprise O₂ or anoxygen-containing compound. The method 300 can proceed to process block308, where at least some of the H₂ produced by the methane conversion istransported (e.g., via permeation of hydrogen ions) across the firstmembrane (e.g., from the first gas flow volume to the second gas flowvolume). The method 300 can then proceed to process block 310, where thetransported H₂ (e.g., in the second gas flow volume) reacts with thefirst sweep gas, for example, to combust the permeated hydrogen withoxygen in the sweep gas to form water and to generate heat that candrive, at least in part, the methane conversion of process block 302.

After process block 310, or if it was not desired to remove hydrogen atdecision block 304, the method 300 can proceed to process block 312,where at least some aromatics can be removed from the first processedstream. For example, aromatics can be removed from (e.g., separatedfrom) the first processed stream via condensing the aromatics into aliquid while other components of the first processed stream remain ingas form. In some embodiments, the aromatics removal can be performed byor within one or more aromatics separation devices, such as any ofdevice 114, device 164, condenser 210, heat exchanger 414, chiller 418,and condenser 422. In some embodiments, the removal of process block 312can produce a second processed stream comprising the C₂ hydrocarbons,any unreacted methane, and any aromatics and/or H₂ not previouslyremoved.

The method 300 can proceed to optional process block 314, where theremoved aromatics (e.g., in liquid form) can be stored, transported,and/or used. For example, the liquid aromatics can be stored in acontainer that is releasably coupled or fixedly coupled to an aromaticsseparation device. The method 300 can proceed to process block 316,where a second sweep gas can be flowed on a first side (e.g., within asecond gas flow volume) of a second membrane (e.g., of a second reactor)opposite to the second processed stream and/or a resulting productstream (e.g., within a first gas flow volume. In some embodiments, thesecond sweep gas can comprise an inert gas. Alternatively oradditionally, in some embodiments, the second sweep gas can comprise O₂or an oxygen-containing compound.

The method 300 can proceed to process block 318, where at least some ofthe H₂ in the second processed stream is removed. For example, the H₂removal can comprise permeation of hydrogen ions through the secondmembrane (e.g., from the first gas flow volume to the second gas flowvolume). In some embodiments, the H₂ removal can be performed by orwithin one or more second reactors, such as any of reactor 126, reactor194, reactor 196, separator 216, reactor 240, reactor 250, and separator430. The H₂ removal can occur at a second temperature, for example, lessthan 1000 K (e.g., ≤800 K). In some embodiments, the removal of processblock 318 can produce a product stream comprising the C₂ hydrocarbons,any unreacted methane, and potentially any H₂ and aromatics notpreviously removed.

Proceeding to decision block 320, if the second sweep gas comprises O₂or an oxygen-containing compound, the method 300 can proceed to processblock 322, where the removed H₂ (e.g., in the second gas flow volume)reacts with the second sweep gas, for example, to combust the permeatedhydrogen with oxygen in the second sweep gas to form water and togenerate heat that can drive, at least in part, the hydrogen permeationof process block 318 and/or the further conversion of methane by thesecond reactor. After process block 322, or if the sweep gas did notcomprise O₂ or an oxygen-containing compound at decision block 320, themethod 300 can proceed to decision block 324.

At decision block 324, it is determined whether the product streamresulting from process block 318 should be re-processed, for example, toincrease methane conversion and/or aromatic yield by subjecting theproduct stream to another cycle of conversion, aromatic separation, andH₂ removal. If re-processing is desired, the method 300 can proceed fromdecision block 324 to process block 326, where the product stream isredirected to be used as the input gas flow stream to the first reactor(or combined with additional methane for use as the input gas flowstream). The method 300 can then return to process block 302. Otherwise,if re-processing is not desired (e.g., if all of the methane has beenconverted or further methane conversion is not possible), the method 300can proceed from decision block 324 to terminal 328, where the methodcan end.

Although illustrated separately, it is contemplated that various processblocks may occur simultaneously or iteratively. For example, the methaneconversion 302, H₂ transport 308 or 318, sweep gas flow 306 or 316,aromatics removal 312, and H₂ reaction 310 or 322 can occursimultaneously despite being illustrated as sequential process blocks.Furthermore, certain process blocks illustrated as occurring afterothers may indeed occur before. For example, a sweep gas flow 306 or 316may be initiated before any initiation of methane conversion 302.Although some of blocks 302-328 of method 300 have been described asbeing performed once, in some embodiments, multiple repetitions of aparticular process block may be employed before proceeding to the nextdecision block or process block. In addition, although blocks 302-328 ofmethod 300 have been separately illustrated and described, in someembodiments, process blocks may be combined and performed together(simultaneously or sequentially). Moreover, although FIG. 3 illustratesa particular order for blocks 302-328, embodiments of the disclosedsubject matter are not limited thereto. Indeed, in certain embodiments,the blocks may occur in a different order than illustrated orsimultaneously with other blocks.

Methane Conversion System

FIG. 4 illustrates a more detailed configuration of an exemplary gasrecycle system 400. In the illustrated example, gas recycle system 400includes a first heat exchanger 404, a feed coupler 406, a DNMC reactor410, a second heat exchanger 414, a chiller 418, a condenser 422, an H₂membrane separator 430, and a valve 436. In operation of system 400, amethane feed stream 402 can be provided to first heat exchanger 404,which can serve to preheat the methane prior to introduction into theDNMC reactor 410. The methane stream can be directed from the first heatexchanger 404 via feed coupler 406 as input flow stream 408 to the DNMCreactor 410. The feed coupler 406 can select between the fresh methanefeed stream or the recycle stream 440 (e.g., product stream 432 viavalve 436 and recycle line 442) for introduction to the DNMC reactor410. Within the DNMC reactor 410, at least some of the methane in inputflow stream 408 can be converted to C₂ hydrocarbons, H₂, and aromatics.A first processed stream 412 comprising any unreacted methane along withthe produced C₂ hydrocarbons, H₂, and aromatics can be output from theDNMC reactor 410.

The first processed stream 412 can be passed through second heatexchanger 414 (e.g., cross-flow heat exchanger) to cool the first stream412. In the illustrated example, the cooling of the first processedstream 412 serves to preheat the second processed stream 426 prior tointroduction to the H₂ separator 430, thereby providing heat recovery.The resulting cooled stream 416 is directed to chiller 418, which servesto further reduce the temperature of the stream in preparation foraromatics separation. The further cooled stream 420 is then directed tocondenser 422, which condenses at least some of the aromatics out of thestream, thereby forming a first output stream 424 of liquid aromaticsand a second processed stream 426 of unreacted methane, H₂, and C₂hydrocarbons (and any aromatics not removed by condenser 422). In someembodiments, the heat exchanger 414, chiller 418, and/or condenser 422may be considered components of an aromatics separation device ormodule.

As noted above, the second processed stream 426 is passed through thesecond heat exchanger 414 for preheating by heat recovery from the firstprocessed stream 412. The resulting preheated stream 428 is directed toH₂ separator 430, where at least some H₂ in the preheated stream 428 canbe removed from the stream, for example, via permeation of hydrogenthrough a membrane, thereby forming a second output stream 434 of purehydrogen (e.g., permeated H₂ product) and a product stream 432 ofunreacted methane and C₂ hydrocarbons (and any aromatics and/or H₂ notpreviously removed). The product stream 432 can be directed forre-processing by the system 400 (e.g., in another cycle) via valve 436and recycle line 442. Alternatively, if re-processing is not desired(e.g., if all methane has been converted and/or further methaneconversion is not possible), then the valve 436 can redirect the productstream 432 to vent 438, for example, for release to atmosphere ordirected for storage, use, or disposal.

System 400 can include additional components beyond those specificallyillustrated in FIG. 4 or described above. For example, system 400 caninclude a controller, a heater or furnace, temperature sensors, flowsensors, flow controllers, pumps, etc.

Computer Implementation

FIG. 5 depicts a generalized example of a suitable computing environment231 in which the described innovations may be implemented, such ascontroller 140, a controller of system 200, a controller of system 400,or method 300. The computing environment 231 is not intended to suggestany limitation as to scope of use or functionality, as the innovationsmay be implemented in diverse general-purpose or special-purposecomputing systems. For example, the computing environment 231 can be anyof a variety of computing devices (e.g., desktop computer, laptopcomputer, server computer, tablet computer, etc.).

With reference to FIG. 5, the computing environment 231 includes one ormore processing units 235, 237 and memory 239, 241. In FIG. 5, thisbasic configuration 251 is included within a dashed line. The processingunits 235, 237 execute computer-executable instructions. A processingunit can be a central processing unit (CPU), processor in anapplication-specific integrated circuit (ASIC) or any other type ofprocessor (e.g., hardware processors, graphics processing units (GPUs),virtual processors, etc.). In a multi-processing system, multipleprocessing units execute computer-executable instructions to increaseprocessing power. For example, FIG. 5 shows a central processing unit235 as well as a graphics processing unit or co-processing unit 237. Thetangible memory 239, 241 may be volatile memory (e.g., registers, cache,RAM), non-volatile memory (e.g., ROM, EEPROM, flash memory, etc.), orsome combination of the two, accessible by the processing unit(s). Thememory 239, 241 stores software 233 implementing one or more innovationsdescribed herein, in the form of computer-executable instructionssuitable for execution by the processing unit(s).

A computing system may have additional features. For example, thecomputing environment 231 includes storage 261, one or more inputdevices 271, one or more output devices 281, and one or morecommunication connections 291. An interconnection mechanism (not shown)such as a bus, controller, or network interconnects the components ofthe computing environment 231. Typically, operating system software (notshown) provides an operating environment for other software executing inthe computing environment 231, and coordinates activities of thecomponents of the computing environment 231.

The tangible storage 261 may be removable or non-removable, and includesmagnetic disks, magnetic tapes or cassettes, CD-ROMs, DVDs, or any othermedium which can be used to store information in a non-transitory way,and which can be accessed within the computing environment 231. Thestorage 261 can store instructions for the software 233 implementing oneor more innovations described herein.

The input device(s) 271 may be a touch input device such as a keyboard,mouse, pen, or trackball, a voice input device, a scanning device, oranother device that provides input to the computing environment 231. Theoutput device(s) 271 may be a display, printer, speaker, CD-writer, oranother device that provides output from computing environment 231.

The communication connection(s) 291 enable communication over acommunication medium to another computing entity. The communicationmedium conveys information such as computer-executable instructions,audio or video input or output, or other data in a modulated datasignal. A modulated data signal is a signal that has one or more of itscharacteristics set or changed in such a manner as to encode informationin the signal. By way of example, and not limitation, communicationmedia can use an electrical, optical, radio-frequency (RF), or anothercarrier.

Any of the disclosed methods can be implemented as computer-executableinstructions stored on one or more computer-readable storage media(e.g., one or more optical media discs, volatile memory components (suchas DRAM or SRAM), or non-volatile memory components (such as flashmemory or hard drives)) and executed on a computer (e.g., anycommercially available computer, including smart phones or other mobiledevices that include computing hardware). The term computer-readablestorage media does not include communication connections, such assignals and carrier waves. Any of the computer-executable instructionsfor implementing the disclosed techniques as well as any data createdand used during implementation of the disclosed embodiments can bestored on one or more computer-readable storage media. Thecomputer-executable instructions can be part of, for example, adedicated software application or a software application that isaccessed or downloaded via a web browser or other software application(such as a remote computing application). Such software can be executed,for example, on a single local computer (e.g., any suitable commerciallyavailable computer) or in a network environment (e.g., via the Internet,a wide-area network, a local-area network, a client-server network (suchas a cloud computing network), or any other such network) using one ormore network computers.

For clarity, only certain selected aspects of the software-basedimplementations are described. Other details that are well known in theart are omitted. For example, it should be understood that the disclosedtechnology is not limited to any specific computer language or program.For instance, aspects of the disclosed technology can be implemented bysoftware written in C++, Java™, Python®, or any other suitable computerlanguage. Likewise, the disclosed technology is not limited to anyparticular computer or type of hardware. Certain details of suitablecomputers and hardware are well known and need not be set forth indetail in this disclosure.

It should also be well understood that any functionality describedherein can be performed, at least in part, by one or more hardware logiccomponents, instead of software. For example, and without limitation,illustrative types of hardware logic components that can be used includeField-programmable Gate Arrays (FPGAs), Program-specific IntegratedCircuits (ASICs), Program-specific Standard Products (ASSPs),System-on-a-chip systems (SOCs), Complex Programmable Logic Devices(CPLDs), etc.

Furthermore, any of the software-based embodiments (comprising, forexample, computer-executable instructions for causing a computer toperform any of the disclosed methods) can be uploaded, downloaded, orremotely accessed through a suitable communication means. Such suitablecommunication means include, for example, the Internet, the World WideWeb, an intranet, software applications, cable (including fiber opticcable), magnetic communications, electromagnetic communications(including RF, microwave, and infrared communications), electroniccommunications, or other such communication means. In any of theabove-described examples and embodiments, provision of a request (e.g.,data request), indication (e.g., data signal), instruction (e.g.,control signal), or any other communication between systems, components,devices, etc. can be by generation and transmission of an appropriateelectrical signal by wired or wireless connections.

Fabricated Examples and Experimental Results Gas Recycle System

To test performance of the gas-recycle system (e.g., having a setupsimilar to that shown in FIG. 2A) with respect to CH₄ conversion andaromatics liquid yield, the DNMC reactor (e.g., reactor 206 in FIG. 2A)and the H₂-permeable membrane separator (e.g., separator 216 in FIG. 2A)were investigated separately with simulated feeds in each cycle. TheDNMC reactor was formed as a fixed-bed quartz tube (e.g., in aconfiguration similar to reactor 230 shown in FIG. 2B, with a 7-mm innerdiameter), in which 0.3750 g Fe@SiO₂ catalyst was packed for methaneactivation. The DNMC reaction was run at atmospheric pressure and 1303 Kin the DNMC reactor, which was controlled using a furnace (Applied TestSystems Series 3210) connected to a controller (Eurotherm TemperatureController, 3216 series). Catalyst temperatures were measured using aK-type thermocouple touching the middle of the catalyst bed on theexternal surface of the quartz tube. In the first cycle of theexperiment, the catalyst was heated to the reaction temperature at aramp rate of 10 K/min in argon (Ar) flow (20 mL/min). Afterward, the gasflow was switched to a reaction gas mixture (20 mL/min, 90% CH₄ and 10%Ar). The product effluents were analyzed on-line by a gas chromatograph(Agilent 6890) containing a (5%-Phenyl)-methylpolysiloxane capillarycolumn (HP-5, 30.0 m×0.25 mm×0.25 μm) connected to a flame-ionizationdetector (FID) and a packed column (ShinCarbon ST Columns, 80/100 mesh,1.83-m) linked to a thermal conductivity detector (TCD). Transfer lineswere maintained at temperatures greater than 473 K by resistive heatingto prevent any condensation. In the second and follow-on reactioncycles, the feed gas was a mixture of H₂, CH₄ and Ca hydrocarbons, whosecomposition and flow rate were determined from the preceding gas flowstream after passing through an aromatics condenser and H₂-separator.The products were analyzed using the same set-up and method as those forthe first cycle of the experiment.

The H₂ separator was an H₂-permeable perovskite ceramic membrane tube(e.g., in a configuration similar to reactor 240 shown in FIG. 2C) thatcomprised a porous SrCe_(0.8)Zr_(0.2)O_(3-δ) support and a thin anddense H₂-permeable SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) film on the poroussupport. The SrCe_(0.8)Zr_(0.2)O_(3-δ) perovskite ceramic is a mixedionic and electronic conductor (MIEC), which means internal electronictransport balances the protonic transport eliminating the need for anexternal electric circuit to enable H₂ separation. The H₂-permeablemembrane tube was prepared by tape casting a SrCe_(0.8)Zr_(0.2)O₃ slurryand rolling end-capped tubular-type supports, followed by colloidalcoating a thin dense SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) layer on theporous support. For the H₂ permeation test of the H₂ separator, the feedside (inside of the membrane tube) was exposed to H₂ diluted to thetested concentration using an Ar tracer. The total flow rate of the feedgas (mixture of H₂ and Ar) was set at 20 mL/min. The sweep side (outsideof the membrane tube) was exposed to 5% N₂/He at 20 mL/min and connectedto the gas chromatograph (Agilent 6890) to quantify the permeated H₂. Inaddition to being a diluent, the Ar in the feed gas was used as a tracerto confirm no trans-membrane leak, which would be indicated by anincrease in Ar signal of the gas chromatograph.

The composition of simulated feed for the DNMC reactor was determinedfrom the stream in the previous cycle after passing through thearomatics liquid condenser and H₂ membrane separator. The composition offeed for the H₂ separator was set the same as the product effluent fromthe DNMC reactor in the same cycle, except that all the aromatics werecondensed and separated from the effluent. Finally, the DNMC reactor andH₂ membrane separator together with an aromatics liquid condenser weresimulated by Aspen Plus® to analyze the system feasibility for CH₄upgrading via the DNMC process.

To understand the membrane H₂ removal capability from the gas recycle,the H₂ permeation fluxes at different temperatures and H₂ partialpressures were determined for a feed flow rate of 20 ml/min (H₂concentration varied, balanced by Ar) and a sweep side flow rate of 20ml/min (5% N₂ balanced by He). FIG. 6A shows that the permeated H₂fluxes as a function of separator temperature and H₂ partial pressure inthe feed side in a single membrane tube. With increasing temperature orH₂ partial pressure, the H₂ flux increased. This is caused by theincreasing ambipolar ionic/electronic conductivity of the MIECSrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) membrane at high temperature or highpartial pressure of H₂ feed. From the flow rate of H₂ feed andH₂-permeation flux, the H₂ removal efficiency, indicated by the ratio ofH₂-permeation flux/H₂ feed, at each condition was evaluated, as shown inFIG. 6B.

The H₂ removal efficiency does not depend on the H₂ concentration in thefeed, but it does depend on the temperature. When the temperature wasincreased from 873 K to 1173 K, the H₂ removal efficiency was increasedfrom 2.67% to 18.60%, equivalent to ˜7 times higher enhancement in H₂removal efficiency at the same H₂ feed condition. Therefore, in somecases, high temperature can be used to remove H₂ more efficiently fromthe H₂-permeable SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) membrane separator.The SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) membrane tube had an inner diameterof 7 mm and a length of 15 cm; however, only 20% (˜3 cm) of the membranetube length was kept at the test temperature due to the thermal gradientof the tubular furnace that supplied heat to the membrane. Thus, theH₂-permeation flux could be significantly higher compared to the presentH₂ flux data of FIGS. 6A-6B if the entire membrane tube was kept at thetest temperature. The H₂ flux can be enhanced significantly if operatedunder a pressure gradient and the total H₂ permeance area increased byuse of a membrane tube bundle.

In the gas recycle system, the H₂ removal stage takes place after theDNMC reaction in the DNMC reactor. The feed to the H₂ separator thuscontains both H₂ and hydrocarbons (e.g., unreacted CH₄ and C₂ products),which feed is slightly different from the feed conditions examined toproduce FIGS. 6A-6B. In order to test the effect of hydrocarbons in theH₂ feed stream on the performance of the H₂ membrane, a controlexperiment was performed by adding 5% CH₄ into the H₂ feed in thepermeation tests, with a feed flow rate feed flow rate of 20 ml/min (50%H₂+5% CH₄, balanced by Ar) and a sweep side flow rate of 20 ml/min (5%N₂ balanced by He). As shown in FIG. 7A-7B, the H₂ permeation fluxincreased with increasing temperature, even in the presence of CH₄ inthe feed. When the temperature was set at 1173 K, a reduction of the H₂flux was observed, which may be caused by carbon deposition on themembrane surface. For example, at high temperatures (e.g., >1100 K), theSrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3-δ) membrane materials could cause CH₄activation to coking species, and thus lead to lower H₂ flux. Thisconfirms the advantages of decoupling the CH₄ reaction and H₂ separationby separating into two different stages (e.g., operation units), sinceDNMC requires high temperature (e.g., >1100 K) to reach appreciable CH₄conversion and C₂₊ yield that would otherwise cause theSrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) membrane performance to deterioratewith respect to H₂ permeation. However, these results apply to the testmembrane, and may not be indicative of all membranes. Moreover, asdescribed above and in further detail below, the use of a sweep gascomprising O₂ or an oxygen-containing compound can prevent, or at leastreduce, coke formation on the membrane and thereby allow the H₂separation membrane to operate at higher temperatures (e.g., higher thanan operating temperature of the catalyst).

It should be noted that the H₂ flux in FIG. 7A was stable up to 1073 K.Moreover, the H₂ permeation flux at 1073 K in the experiments underlyingFIG. 7A was the same as that of 50% H₂ feed (balanced by Ar and in theabsence of CH₄ hydrocarbon) in the experiments underlying FIG. 6A.Therefore, for the remaining H₂ permeation tests for the gas recyclesystem, the temperature of the H₂ membrane separator was set at 1073 K,and the H₂ removal efficiency at this condition was 10% whetherhydrocarbons were present or not in the feed to the H₂ membrane.

To examine the performance of the CH₄ reaction in the gas recyclesystem, the feed and exit effluent compositions of the DNMC reactor weremeasured, which compositions can change depending on the recycle numberof the process. Therefore, the gas compositions and gas flow rates ineach cycle were measured before the overall performance of the gasrecycle system was analyzed. In DNMC, the gas feed was started (i.e., inthe first cycle) with 90% CH₄ mixed with 10% Ar internal standard at aflow rate of 20 mL/min, as shown in FIG. 8A. At the reaction temperatureof 1303 K, the DNMC reactor achieved 14.0% CH₄ conversion, 5.4% C₂product yield, and 8.5% aromatics yield, as shown in FIG. 9A. The DNMCover Fe@SiO₂ catalyst showed stable CH₄ conversion and productselectivity. Assuming all aromatics products are removed via a condenserafter the DNMC reactor, the feed to the H₂ separator would be comprisedof unreacted CH₄, C₂ hydrocarbons, and H₂ gas in the effluents of themethane reactor. On the basis of the H₂ removal efficiency of the H₂separator in FIGS. 6A-6B, the effluents after the H₂ separator are thenfed to the DNMC reactor in the subsequent cycle.

Since ethylene (C₂H₄) is the predominant C₂ product, as shown in FIG.9A, the Ca species in the effluent gases after the H₂ separator wereassumed to be C₂H₄. Therefore, the gas feed for the subsequent cyclesfor the DNMC reactor in the gas recycle system was composed of CH₄,C₂H₄, and H₂. For example, after the DNMC reactor (14.0% CH₄ conversion)and H₂ separator (10.0% H₂ removal efficiency) in the first cycle in theexperimental setup, the feed to DNMC reactor was comprised of a gasmixture of CH₄/C₂H₄/H₂ at a molar ratio of 1.00/0.03/0.20 and at a flowrate of 19.13 mL/min. Following the same analysis described above, FIG.8A shows the composition and flow rate of the feed gas to the DNMCreactor at each cycle number at 10% H₂ removal efficiency by the H₂separator in each cycle. It can be seen that the flow rate of gasmixture CH₄/C₂H₄/H₂ increased and then decreased with the cycle number.In the third cycle, the gas mixture flow rate was increased to themaxima (19.87 mL/min, equivalent to a 9.32% increase compared to the18.00 mL/min flow rate in the first cycle). Up to the ninth cycle, thegas mixture flow rate was still 1.00% higher than the feed flow rate inthe first cycle. The gas composition in the feed to the DNMC reactorchanges with the cycle number. With increasing reaction cycle, the CH₄composition was monotonically decreased, H₂ concentration had anopposite trend, and C₂H₄ concentration (˜2.40%) was nearly constantindependent of the reaction cycle. At the ninth reaction cycle, the CH₄flow rate dropped to 51.80% of the CH₄ feed in the first cycle, whichindicates the rest of CH₄ (˜49.20%) has been converted to aromaticsliquids after nine reaction cycles.

Although the single H₂ separator tube in the experiments exhibited only10% H₂ removal efficiency, further enhancement in H₂ removal efficiencycan be achieved, for example, as described elsewhere herein. To test howH₂ removal efficiency affects the performance of the DNMC reactor in thegas recycle system, the composition and flow rate of CH₄/C₂H₄/H₂ feedmixture in each cycle was analyzed by assuming the H₂ separator reaches40%, 70%, and 100% H₂ removal efficiency, respectively. As shown inFIGS. 8B-8D, the total gas flow rate decreased with increasing H₂removal efficiency in the H₂ separator unit at the same cycle. Withincreasing H₂ removal efficiency at the same cycle, the CH₄concentration in the CH₄/C₂H₄/H₂ feed mixture increases and thus H₂concentration decreases. The C₂H₄ concentration increased to 2.84%,3.25% and 3.53% when the H₂ membrane separator had removal efficienciesof 40%, 70% and 100%, respectively. Compared to the number of cycles inFIG. 8A, —50% CH₄ conversion can be achieved in a smaller number ofcycles for the higher H₂ removal efficiencies of FIGS. 8B-8D. However,in contrast to FIG. 8A, the feed flow rate of FIGS. 8B-8D decrease withincreasing cycle number. Product yield is limited by H₂ separation flux.By increasing the temperature, the H₂ flux will increase, and thusproduct yield will also increase for each cycle. As noted above, suchhigher temperature operations can be enabled by the use of a sweep gascomprising O₂ or an oxygen-containing compound, in particular, whereback-diffusion of oxygen avoids, or at least reduces, coking of themembrane.

The feed stream with the same composition and flow rate as determinedabove for the DNMC reactor in Section 3.2.1 was tested over the Fe@SiO₂catalyst in the DNMC reactor. FIGS. 9A-9D show the CH₄ conversion andproduct selectivity at different cycle numbers and different H₂ removalefficiency conditions. At 10% H₂ removal efficiency in each cycle, CH₄conversion decreased as the cycle number increased, as shown in FIG. 9A.Two different phenomena can govern the effects of the CH₄ conversiontrend. From the second cycle onward, C₂ and H₂ were included in the feedgas as a result of the DNMC reaction from the previous cycle, which willaffect the equilibrium CH₄ conversion. According to the Le Châtelier'sprinciple, the inclusion of H₂ will shift the equilibrium to thereactant side lowering the CH₄ conversion. From the feed composition ateach cycle (FIG. 8A), an increase in H₂ concentration of the feed gaswas observed, which agreed with the decreasing trend of the CH₄conversion. In addition, from FIG. 8A, the total flow rates of thesubsequent cycles were slightly higher than the first cycle, which meanshigher space velocities for the DNMC reaction that in turn leads tolower CH₄ conversion. As for the selectivity, the product selectivityshifted toward lighter products as the cycle number increased, due tothe addition of H₂ in the feed of the subsequent cycles, which shiftsthe overall equilibrium of the DNMC reaction to lighter hydrocarbons.

Based on the feed flow rate and composition analysis of FIGS. 8B-8D, thearomatics liquid production from DMNC reaction can be enhancedsignificantly if the H₂ removal efficiency is further improved. Again,the H₂ flux can be increased by increasing the temperature, and suchhigher temperature operations can be enabled by the use of a sweep gascomprising O₂ or an oxygen-containing compound to avoid, or at leastreduce, coking of the membrane. FIG. 8B shows the performance of the CH₄reactor at 40% H₂ removal in the H₂ separator in each cycle. First, theH₂ content leftover in the feed of the subsequent cycles (FIG. 8B) wereless compared to the case of the 10% H₂ removal. However, the amount ofH₂ in this feed stream (up to ˜30%) should still lead to lower CH₄conversion, and shift product selectivity towards lighter hydrocarbons.Second, the overall gas feed flow rates dropped significantly (up to˜50%), which led to lower space velocities for the DNMC reaction. Lowerspace velocities of DNMC should lead to higher CH₄ conversion and shiftthe product selectivity toward larger hydrocarbons. These two effectsseemed to balance each other out at the case of 40% H₂ removalefficiency. As shown in FIG. 9B, a relatively stable CH₄ conversion andproduct selectivity were achieved in this case in each reaction cycle.At significant levels of H₂ removal efficiencies of 70% (FIG. 9C) and100% (FIG. 9D), the effects of lowering DNMC reaction space velocitiesdominate over the inclusion of H₂ in the subsequent cycles. Therefore,the CH₄ conversion increased as the cycle number increased for 70% H₂removal efficiency and 100% H₂ removal efficiency. In addition, theproduct selectivity shifted toward larger aromatic hydrocarbons(liquid). Too high of a reduction in the DMNC space velocities led tocoke formation, which was demonstrated at the fourth cycle of thereaction with 100% H₂ removal in FIG. 9D.

Based on the above-described separate performance evaluations of theDNMC reactor and the H₂ membrane separator, the overall performance ofthe gas recycle system (integrating the DNMC reactor and H₂ membraneseparator) can be evaluated to understand the overall CH₄ conversion,product selectivity, and yields. In particular, FIGS. 10A-10D show theoverall CH₄ conversion and product selectivity at different levels of H₂removal efficiency, with H₂ separation performed at 1073 K using asingle H₂ permeable membrane). As is apparent from FIGS. 10A-10D, all ofthe H₂ removal levels showed similar trends in methane conversion andproduct selectivity. As the number of cycles increased, the overall CH₄conversion increased, with the overall CH₄ conversion reaching over 70%and potentially higher. The product selectivity also exhibited the sametrend. As the number of cycles increased, the product selectivityshifted toward larger aromatic hydrocarbons (liquid), with the liquidportion of the products reaching over 95%. As the amount of H₂ removalincreases, it takes a fewer number of cycles to reach the same level ofCH₄ conversion, since the equilibrium conversion is enhanced by shiftingtoward the product side according to the Le Châtelier's principle.

FIG. 11 shows the overall aromatics yields at different H₂ removalefficiencies. In the experimental setup comprising the DNMC reactor andthe H₂ separator using a single SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ)membrane tube, the yield of aromatics was >40% after nine cycles for 10%H₂ removal, as shown in FIG. 11. With higher H₂ removal efficiencies,the system has the potential to reach even higher aromatics yield. Forexample, at 40% H₂ removal, the aromatics yield can reach over 70%.Comparing the aromatics yields of the system with different H₂ removalefficiencies clearly shows that the higher H₂ removal capability of theH₂ separator leads to higher aromatics yield, since removal of H shiftsthe equilibrium to the product side, which increases the equilibrium CH₄conversion.

Autothermal Membrane Reactor Operation

To test performance of an autothermal reactor (e.g., having a setupsimilar to heat-exchange reactor 290 in FIG. 2E), a DNMC reaction wasrun at atmospheric pressure in both a fixed-bed reactor (without H₂permeation, similar to reactor 230 in FIG. 2B) and a tubular membranereactor (with H₂ permeation). In the experiments, 0.375 g of Fe@SiO₂catalyst was loaded at the center of the reactor and then heated to thedesired temperature in pure Ar at the rate of 20 mL/min. After thereaction gas mixture (90% CH₄ and 10% Ar) was introduced, the reactionwas run at a temperature range of 1253-1303 K and at a feed gas spacevelocity of 3200 mL/(g-h). During the reaction in the heat-exchangeH₂-permeable membrane reactor, a mixture of 20 mol % O₂ balanced by Hewas introduced as the sweep gas. The concentration of O₂ in the sweepgas is similar to that of air, but He was used instead of N₂ in theexperiments due to the high-sensitivity of flame ionization detector(FID) and thermal conductivity detector (TCD) caused by He carrier gasin the gas chromatography. The CH₄ reactant was introduced through theinner tube in the top center section of the heat-exchange H₂-permeablereactor and the C₂₊ products exited through the catalyst bed in thereactor. The outer annular region of the membrane reactor was exposed tothe O₂/He sweep gas to carry and react with the permeated H₂ away fromthe reactor system. The effluent gases from the feed side and the sweepside were analyzed on-line by a gas chromatographer (Agilent 6890).Transfer lines were maintained at temperatures greater than 473 K byresistive heating to prevent product condensation.

FIGS. 12A-12B show the effects of different sweep gas on the DNMCreaction, which is on the inside of the heat-exchange H₂-permeablemembrane reactor. As a reference, a control experiment was performedwithin a fixed-bed reactor that is made of a quartz tube. FIG. 12A showsan increase in CH₄ conversion when comparing the fixed-bed reactor(without H₂ removal) to the membrane rector with or without the sweepgas flow. Without a sweep gas, CH₄ conversion in the membrane reactorwas slightly higher compared to the fixed-bed reactor, which is due toactivation of methane by the membrane material itself. A controlexperiment was done by placing the SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ)oxide powder in the fixed-bed reactor, which showed the methaneconversion with plenty (>50% selectivity) of coke formation at thetested DNMC reaction conditions. The higher coke selectivity of DNMC inthe membrane reactor than that of the fixed bed reactor in the absenceof any sweep gas, as shown in FIG. 12B, also confirmed the activation ofmethane into coke by the membrane material itself. By using the sweepgas to simultaneously remove the permeated H₂ from the membrane reactor,the reaction is shifted to the product side which increases the CH₄conversion in accordance with Le Châtelier's principle. Therefore, theH₂-permeable membrane reactor in the presence of He-only or O₂/He sweepgas showed higher CH₄ conversion than that of the fixed-bed reactor, asshown in FIG. 12A. When O₂ is present in the sweep gas, the CH₄conversion was slightly decreased, which can be attributed to thesacrifice of H₂ permeation by back diffusion of O₂ in theSrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) membrane since CO product is formed inthis condition. The product yields, especially the aromatics yield, inall the DNMC test conditions follow the same trend of CH₄ conversions.

FIG. 12B compares the product selectivity in DNMC reaction in thefixed-bed and membrane reactors with different sweep gas conditions.When compared to the fixed-bed reactor, the membrane reactor withoutsweep gas flow showed a slight decrease in C₂ selectivity, and a slightincrease in both benzene and coke selectivities. Adding an He-only sweepgas in the membrane reactor further shifted the C₂ and benzene/coke tolower and higher selectivities, respectively. When O₂ was present in thesweep gas, the C₂ selectivity kept similar to that of He only sweep gascondition, but the benzene and coke selectivities were decreased. Moreimportantly, coke was not produced, and CO product was formed instead.Therefore, the disclosed membrane reactor systems can have bothlong-term stability and low-carbon efficiency constraints, for example,by completely eliminating the coke formation and having >90% carbonefficiency.

The formation of CO in the DNMC reaction in the heat-exchangeH₂-permeable membrane reactor can be attributed to the fact that theMIEC SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) membrane material can co-permeateboth H₂ and O₂ gases (e.g., via respective ions thereof). In particular,the back diffusion of O₂ from the O₂/He sweep gas across the membraneinto the DNMC reaction volume can oxidize any carbon depositionresulting from the DNMC reaction into CO, and thereby eliminate, or atleast reduce, coke formation. Concurrently, when O₂ is used in the sweepgas, the ambipolar conductivity of the SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3-δ)membrane may decrease, which can lead to a decrease in the rate of H₂permeation, and therefore lower CH₄ conversion. On the other hand, whenpermeated H₂ is readily combusted by O₂ in the sweep gas, there is anincrease in the H₂ partial pressure difference across both sides of themembrane, which can lead to an increase in H₂ permeation, and thereforehigher CH₄ conversion. As shown in FIG. 12A, a slight decrease inoverall CH₄ conversion is observed when switching from He-only to O₂/Hesweep gas, suggesting that the trade-off between effects of the decreasein ambipolar conductivity and effects of the increase in H₂ partialpressure gradient is dominated by the former property in the testedconditions.

The effects of reaction temperature on methane conversion, product yieldand selectivity, H₂ production and permeation, as well as heatrequirement for DNMC and heat release from combustion of H₂ permeatewere studied, and the results are summarized in FIGS. 13A-13D. Anincrease in CH₄ conversion with increasing temperature is observed inFIG. 13A, which is due to the endothermic nature of the DNMC reaction.The product yields also increases with increasing reaction temperature.In particular, the aromatics yield increases significantly and C₂hydrocarbon yield increased moderately, while CO only exhibited a slightincrease, as shown in FIG. 13A. FIG. 13B shows the product selectivityversus reaction temperature in DNMC in the heat-exchange H₂-permeablemembrane reactor. As the temperature is increased, the productselectivity shifts from smaller C₂ products to aromatics. This trendcorresponds to the thermodynamic nature of the DNMC reaction, whichmeans that high reaction temperature and high CH₄ conversion favor heavyproduct formation. The CO selectivity, however, exhibited an oppositetrend with increasing reaction temperature. Regardless, there was nocoke formation at any of the tested temperature conditions.

FIG. 13C shows the rates of H₂ generation from DNMC and H₂ permeationthrough the MIEC membrane at the three tested temperatures (1253 K, 1273K, and 1303 K). The H₂ production rate was calculated from the carbonbalance in the DNMC reaction, while the H₂ permeation rate was evaluatedon the basis of hydrogen balance in both DNMC and H₂ combustionreactions. When the reaction temperature was increased (e.g., from 1253K, or 1273 K, to 1303 K), a significant increase in the H₂ productionrate was observed. The H₂ permeation rate through the membrane increasedslightly with the reaction temperature. Therefore, the percentages of H₂removal were calculated to be 36.9%, 27.9% and 17.5% at 1253 K, 1273 K,and 1303 K, respectively. Since O₂ content is in excess compared to H₂permeate in the sweep gas side, the H₂ permeate was completely consumedby the combustion reaction outside of the MIEC membrane tube.

From the oxygen balance in the membrane reactor, the O₂ conversions werefound to be 7.7%, 9.3% and 10.6% at the reaction temperatures of 1253 K,1273 K, and 1303 K, in sequence. Thus, the O₂ residue in the sweep gasdecreases with increasing reaction temperature. As noted above, the backdiffusion of O₂ from the O₂/He sweep gas (through the MIEC membrane intothe DNMC reaction volume) was responsible for CO formation. According tothe Wagner Equation, the O₂ permeation should increase with temperature,but a decrease in O₂ concentration on the sweep side of the membrane canlead to lower O₂ permeation. The interaction between these two opposingeffects can offset each other, such that the CO yield at differentreaction temperatures remains about the same, as shown in FIG. 13A.

As shown in FIG. 13B, the significant increase in C₂₊ yields resulted ina decrease in CO selectivity with increasing reaction temperature. Thematerial balance analyses in FIGS. 13A-13C, together with heat ofreactions for both DNMC and H₂ combustion, enables the autothermalityanalysis for the H₂-permeable membrane reactor. FIG. 13D compares theheat requirement for DMNC and the heat release from the combustion of H₂permeate at 1253 K, 1273 K, and 1303 K. At the lower temperature (1253K), the heat released from H₂ combustion was higher than that of heatrequirement for DNMC reaction. At the middle temperature (1273 K), theheat released from H₂ combustion and the heat requirement for DNMCreaction are almost identical to each other. When the reactiontemperature is increased to the higher temperature (1303 K), the heatreleased from H₂ combustion is not enough to match the heat requirementfor the DNMC reaction, which is caused by the low H₂ removal (17.5%).Overall, the results confirm that the autothermality of DNMC in theheat-exchange H₂-permeable membrane reactor can be achieved at ˜30% H₂removal from the DNMC reaction. Ambipolar conductivity of the MIECSrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3-δ) membrane can be improved in order toreach the balance between the endothermic heat requirement of DNMCreaction and the heat released from combustion of H₂ permeate at highreaction temperatures.

To study tunability of autothermality of DNMC in the heat-exchangeH₂-permeable membrane reactor, the effects of O₂ sweep gas flow rates onthe DNMC reaction and heat generation/consumption in both DNMC and H₂combustion reactions were measured. FIGS. 14A-14I show the CH₄conversion, product yield and selectivity, and heat release by H₂combustion and heat requirement for DNMC as a function of sweep gas flowrate (10-80 mL/min) at 1253 K, 1273K, and 1303 K. At 1303 K, theincreasing O₂ sweep gas flow rate led to a slight decrease in CH₄conversion and thus C₂₊ product yield, as shown in FIG. 14A. When thereaction temperature is lower (e.g., 1253 K), CH₄ conversion increasedslightly and then decreased with increasing sweep gas flow rate, asshown in FIG. 14G. At the reaction temperature of 1273 K, the changingtrend of CH₄ conversion and C₂₊ product yield with respect to sweep gasflow rate stayed between that of 1253 and 1303 K, which showed adecrease and then a plateau with increasing O₂ sweep flow rate, as shownin FIG. 14D. Overall, there was no obvious change in the CH₄ conversionsand C₂₊ product yields with respect to the sweep gas flow rates. FIGS.14B, 14E, and 14H show the product selectivity as a function of O₂ sweepgas flow rate at 1303 K, 1273 K, and 1253 K, respectively. Withincreasing O₂ sweep gas flow rate, the CO selectivity increased,accompanied with reduction in aromatics selectivity.

The increase in O₂ sweep gas flow rate could cause two competingphenomena in the heat-exchange membrane reactor that can influence theCH₄ conversion. The increasing O₂ sweep flow rate can increase O₂exposure to the SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3-δ) membrane per unittime, and thus the surface exchange rate of O₂ with the membranematerial. This can have a negative effect on the ambipolar conductivityof the membrane, which leads to a decrease in the flux of the H₂permeation and therefore lowers CH₄ conversion. At the same time, thepermeated H₂ is combusted at a higher rate with increasing O₂ sweep gasflow rate, which leads to an increase in H₂ partial pressure gradientand an increase in CH₄ conversion. As indicated by the decrease in CH₄conversion with increasing O₂ sweep flow rate, a decrease in theambipolar conductivity of the SrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) membranedominates the overall performance, even though it is somewhat offset bythe increase in H₂ partial pressure gradient.

FIGS. 14C, 14F, and 14I compare the heat requirement for DMNC and heatrelease from the combustion of permeated H₂ at different O₂ sweep flowrates. At 1303 K, the heat released from the combustion of permeated H₂was not enough to supply heat for the DNMC reaction, even with high O₂sweep gas flow rate, as shown in FIG. 14C. This was due to theinsufficient H₂ flux across the membrane, compared to the H₂ generationfrom the DNMC reaction. In addition, the heat requirement for DNMCdecreased and the heat release from H₂ combustions increased as the O₂sweep flow rate increased from 10 to 40 mL/min. The former case wascaused by the decreasing CH₄ conversion, while the latter was caused bythe high H₂ permeation and combustion. When the O₂ sweep flow rate wasvery high (e.g., 80 mL/min), the heat release from H₂ combustion wasdecreased slightly, which may have been caused by the deterioration ofambipolar conductivity of the membrane under very high O₂ sweep gas flowconditions.

At 1273 K, the overall CH₄ conversion decreased when compared to 1303 K,and therefore the heat requirement from the DNMC side was reduced, asshown in FIG. 14F. The results showed that the heat released from thecombustion of permeated H₂ matched the required heat input for the DNMCreaction, which meant that autothermal operation was achieved in all theO₂ sweep gas flow rates. When the reaction temperature was reduced to1253 K, the heat release from combustion of H₂ permeate overcomes thatrequired by DNMC reaction, as shown in FIG. 14I. This phenomenon wasmaintained regardless of the O₂ sweep flow rate. Overall, the resultsshow that the flow rate of O₂ sweep gas does not modulate theautothermality of DNMC.

The stability of the heat-exchange H₂-permeable membrane reactor wastested by running DNMC at 1303 K for 50 hours while flowing the O₂ sweepgas, the results of which stability test are shown in FIG. 15A. Aninduction period of ˜5 hours was observed in the initial state of thereaction where the aromatics yield was low. After the induction period,the membrane reactor reached a stable performance, where the CH₄conversion remained at ˜18.0%. The product yields of C₂ products,aromatics, C₂₊ products (i.e., C₂+aromatics), and CO were stable atapproximately 4.6%, 11.9%, 16.5%, and 0.7%, respectively. The COformation, due to back diffusion of oxygen from the sweep side into theDNMC reaction volume, limited coke accumulation in the membrane reactor,and thus maintained the long-term stability of the reactor. The COselectivity was approximately 3.3%, and almost 97% carbon conversionefficiency was obtained in the studied conditions.

FIG. 15B shows the heat requirement for DNMC and heat release fromcombustion of H₂ permeate on opposite sides of theSrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3δ) membrane. After the induction period,the heat requirement for DNMC and the heat release from the H₂combustion were stable throughout the long-term test. The selectedreaction temperature (1303 K) in the long-term stability tests was theharshest condition among all the tested reaction conditions, and thus itis expected that the heat-exchange H₂-permeable membrane reactor canmaintain autothermality and have stable performance under lower reactiontemperatures (e.g., 1253 K or 1273 K).

To further examine feasibility of the autothermal operation of DNMCprocess in the heat-exchange membrane reactor at an industrial scale,Aspen Plus® was used to simulate the scenario on the basis of theexperimental results at 1273 K, where ˜30% of H₂ was removed from theDNMC reaction and was completely combusted by the sweep gas (e.g., O₂ inair). At 1273 K, the total heat requirement for the entire process wasfound to be 2,360,715 kJ/h, while the total heat released was found tobe −2,408,618 kJ/h, resulting in a net −47,903 kJ/h of heat releasedoverall. Furthermore, the simulation results of the heat exchangerdemonstrate that the heat released from the cooling of both the DNMCproduct stream and the H₂ combustion stream are sufficient to heat theDNMC and H₂ combustion feed streams. The results of the Aspen Plus®simulation further showed that autothermal operation was feasible forthe scaled-up DNMC process, indicating the potential for autothermalityof DNMC operation under realistic conditions.

To supplement the Aspen Plus® simulation, additional COMSOL simulationswere performed to understand the effects of gas flow velocities, CH₄concentration, and temperature profiles of the membrane reactor on theperformance of the DNMC reaction. In particular, a COMSOL 2Daxisymmetric model of the DNMC reaction, coupled with theSrCe_(0.7)Zr_(0.2)Eu_(0.1)O_(3-δ) membrane reactor with O₂ sweep gas,was generated. The COMSOL simulation was conducted on the catalyst bedregion of the membrane reactor for both DNMC channel and hydrogencombustion side. The gas velocities fell near the reactor wall and themembrane walls (i.e., in both DNMC and H₂ combustion sides) due to theno-slip boundary conditions. The CH₄ concentration fell from theentrance to the exit of the catalyst bed due to the reaction of CH₄ inthe DNMC catalyst bed. From the center to the wall of the membranereactor, the CH₄ concentration decreased gradually. The rapid decreasein CH₄ concentration close to the reactor wall was due to the low gasvelocity, which in turn limited the convection of fresh reactant to thatwall. FIG. 16 shows the temperature profile of the DNMC catalyst bedinside the membrane reactor. As shown in FIG. 16, changes in temperaturewere insignificant, as the heat provided by the furnace played animportant role in maintaining the reactor at 1273 K, and the maximumtemperature difference was only ˜0.2 K. The outer wall temperature wasset to 1273 K, which strongly affected the temperature distribution. Theheat generated from H₂ combustion was insignificant, and it did notaffect the overall temperature. However, it was enough energy to balanceout the heat required by the DNMC reaction.

CONCLUSION

Any of the features illustrated or described herein, for example, withrespect to FIGS. 1A-16, can be combined with any other featureillustrated or described herein, for example, with respect to FIGS.1A-16 to provide systems, devices, methods, and embodiments nototherwise illustrated or specifically described herein. For example, anyof the reactors illustrated in FIGS. 2C-2E (e.g., modified to include anappropriate catalyst) and/or any of the reactors illustrated in FIGS. 2Band 2E can be used for reactor 107 in FIG. 1A, reactor 106 in FIG. 1B,any of reactors 107, 156 in FIG. 1C, any of reactors 184, 186 in FIG.1D, reactor 206 in FIG. 2A, or reactor 410 in FIG. 4. In anotherexample, any of the reactors illustrated in FIGS. 2C-2E (e.g., with orwithout an appropriate catalyst) can be used for reactor 127 in FIG. 1A,reactor 126 in FIG. 1B, any of reactors 156, 127 in FIG. 1C, any ofreactors 194, 196 in FIG. 1D, separator 216 in FIG. 2A, or membraneseparator 430 in FIG. 4.

All features described herein are independent of one another and, exceptwhere structurally impossible, can be used in combination with any otherfeature described herein. In view of the many possible embodiments towhich the principles of the disclosed technology may be applied, itshould be recognized that the illustrated embodiments are only examplesand should not be taken as limiting the scope of the disclosedtechnology. Rather, the scope is defined by the following claims. Wetherefore claim all that comes within the scope and spirit of theseclaims.

1. A methane conversion system comprising: a first membrane reactorcomprising a first gas flow volume, a second gas flow volume, and afirst membrane separating the first gas flow volume from the second gasflow volume, the first gas flow volume having a first catalyst therein;a first gas supply coupled to the second gas flow volume and constructedto provide a first sweep gas to the second gas flow volume, the firstsweep gas comprising O₂ or an oxygen-containing compound; an aromaticsseparation device connected to receive a first processed stream from thefirst gas flow volume; a second membrane reactor comprising a third gasflow volume, a fourth gas flow volume, and a second membrane separatingthe third gas flow volume from the fourth gas flow volume, the third gasflow volume having a second catalyst therein; a second gas supplycoupled to the fourth gas flow volume and constructed to provide asecond sweep gas to the fourth gas flow volume, the second sweep gascomprising O₂ or an oxygen-containing compound; and a recycle linecomprising one or more fluid conduits, the first gas flow volume of thefirst membrane reactor being connected to receive a recycle stream fromthe third gas flow volume of the second membrane reactor via the recycleline, wherein the first reactor is constructed to convert at least someCH₄ in an input gas flow stream provided to the first gas flow volume ofthe first reactor, so as to provide a first processed stream and suchthat a quantity of CH₄ in the first processed stream is less than thatin the input gas flow stream, the first processed stream comprising CH₄,C₂ hydrocarbons, and aromatics, the C₂ hydrocarbons are acetylene(C₂H₂), ethylene (C₂H₄), ethane (C₂H₆), or any combination of theforegoing, the aromatics are benzene (C₆H₆), toluene (C₇H₈), naphthalene(C₁₀H₈), or any combination of the foregoing, the first membrane isconstructed such that at least some H₂ is removed from the first gasflow volume by hydrogen ions permeating through the first membrane intothe second gas flow volume and such that oxygen ions permeate throughthe first membrane from the second gas flow volume into the first gasflow volume so as to reduce coking of the first membrane, the firstreactor is constructed for autothermal operation via an exothermicreaction between the permeated hydrogen in the second gas flow volumeand the O₂ or oxygen-containing compound in the second gas flow volume,the aromatics separation device is constructed to remove at least somearomatics from the received first processed stream, so as to provide asecond processed stream comprising CH₄ and C₂ hydrocarbons, and toprovide a first output stream comprising the removed at least somearomatics, a quantity of the aromatics in the second processed streambeing less than in the first processed stream, the second reactor isconstructed to convert at least some CH₄ in the second processed streamprovided to the third gas flow volume of the second reactor, so as toprovide a recycle stream to the recycle line and such that a quantity ofCH₄ in the recycle stream is less than in the second processed stream,the second processed stream comprising C₂ hydrocarbons and aromatics,the second membrane is constructed such that at least some H₂ is removedfrom the third gas flow volume by hydrogen ions permeating through thesecond membrane into the fourth gas flow volume and such that oxygenions permeate through the second membrane from the fourth gas flowvolume into the third gas flow volume so as to reduce coking of thesecond membrane, and the second reactor is constructed for autothermaloperation via an exothermic reaction between the permeated hydrogen inthe fourth gas flow volume and the O₂ or oxygen-containing compound inthe fourth gas flow volume.
 2. The methane conversion system of claim 1,wherein the first catalyst, the second catalyst, or both the first andsecond catalysts comprise Fe@SiO₂.
 3. The methane conversion system ofclaim 1, wherein the first membrane, the second membrane, or both thefirst and second membranes comprise a perovskite-type oxide having aformula of M′Ce_(1-x-y)Zr_(x)M″_(y)O_(3-δ), where: M′ is a least one ofSr and Ba; M″ is at least one of Ti, V, Cr, Mn, Fe, Co Ni, Cu, Nb, Mo,W, Pr, Nd, Pm, Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm and Yb; x is between 0.01and 0.2, inclusive; and y is between 0.01 and 0.3, inclusive.
 4. Amethane conversion system comprising: a first reactor having an inletand an outlet; an aromatics separation device having an inlet, a firstoutlet, and a second outlet, the inlet of the separation deviceconnected to receive a first processed stream from the outlet of thefirst reactor; a second reactor having a first gas flow volume, a secondgas flow volume, and a membrane separating the first gas flow volumefrom the second gas flow volume, the first gas flow volume beingconnected to receive a second processed stream from the first outlet ofthe aromatics separation device; and a recycle line comprising one ormore fluid conduits, the inlet of the first reactor being connected toreceive a recycle stream from the first gas flow volume via the recycleline, wherein the first reactor is constructed to convert at least someCH₄ in an input gas flow stream provided to the inlet of the firstreactor, so as to provide to the outlet of the first reactor the firstprocessed stream and such that a quantity of CH₄ in the first processedstream is less than that in the input gas flow stream, the firstprocessed stream comprising CH₄, C₂ hydrocarbons, H₂, and aromatics, theC₂ hydrocarbons are acetylene (CAL), ethylene (C₂H₄), ethane (C₂H₆), orany combination of the foregoing, the aromatics are benzene (C₆H₆),toluene (C₇H₈), naphthalene (C₁₀H₈), or any combination of theforegoing, the aromatics separation device is constructed to remove atleast some aromatics from the first processed stream provided to theinlet of the aromatics separation device, so as to provide to the firstoutlet of the aromatics separation device a second processed streamcomprising CH₄, C₂ hydrocarbons, and H₂, and to provide to the secondoutlet of the aromatics separation device a first output streamcomprising the removed at least some aromatics, a quantity of thearomatics in the second processed stream being less than in the firstprocessed stream, and the second reactor is constructed to remove atleast some H₂ from the second processed stream, which is provided to thefirst gas flow volume, into the second gas flow volume via the membrane,so as to provide to the recycle line a recycle stream comprising CH₄ andCa hydrocarbons, a quantity of the H₂ in the recycle stream being lessthan that in the second processed stream.
 5. The methane conversionsystem of claim 4, further comprising: a gas supply coupled to an inletof the second gas flow volume and constructed to flow a sweep gasthrough the second gas flow volume, the sweep gas supplied by the gassupply comprising O₂ or an oxygen-containing compound, wherein themembrane is constructed such that the at least some H₂ is removed byhydrogen ions permeating through the membrane from the first gas flowvolume into the second gas flow volume and such that oxygen ionspermeate through the membrane from the second gas flow volume into thefirst gas flow volume.
 6. The methane conversion system of claim 5,wherein the second reactor is constructed for autothermal operation viaan exothermic reaction between the permeated hydrogen in the second gasflow volume and the O₂ or the oxygen-containing compound in the secondgas flow volume to form water.
 7. The methane conversion system of claim4, further comprising a storage container in fluid communication withthe second outlet of the aromatics separation device and constructed tostore the removed at least some aromatics therein.
 8. The methaneconversion system of claim 4, wherein the first reactor comprises acatalyst of Fe@SiO₂.
 9. The methane conversion system of claim 4,wherein: the membrane comprises a perovskite-type oxide having a formulaof M′Ce_(1-x-y)Zr_(x)M″_(y)O_(3-δ), where: M′ is a least one of Sr andBa; M″ is at least one of Ti, V, Cr, Mn, Fe, Co Ni, Cu, Nb, Mo, W, Pr,Nd, Pm, Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm and Yb; x is between 0.01 and0.2, inclusive; and y is between 0.01 and 0.3, inclusive; and themembrane is provided on a porous support comprising a perovskite-typeoxide having a formula of M′Ce_(1-z)Zr_(z)O_(3-δ), where M′ is Sr or Ba,and z is between 0.01 and 0.3, inclusive.
 10. The methane conversionsystem of claim 4, wherein the aromatics separation device comprises acondenser constructed to liquefy the at least some aromatics.
 11. Amethod comprising: (a) converting, via a first reactor, at least someCH₄ in an input gas flow stream into C₂ hydrocarbons, H₂, and aromatics,thereby providing a first processed stream comprising CH₄, C₂hydrocarbons, H₂, and aromatics, a quantity of CH₄ in the firstprocessed stream being less than that in the input gas flow stream, theC₂ hydrocarbons being acetylene (CAL), ethylene (C₂H₄), ethane (C₂H₆),or any combination of the foregoing, the aromatics being benzene (C₆H₆),toluene (C₇H₈), naphthalene (C₁₀H₈), or any combination of theforegoing; (b) removing, via an aromatics separation device downstreamof the first reactor, at least some aromatics from the first processedstream, thereby providing a first output stream comprising the removedat least some aromatics and a second processed stream comprising CH₄, C₂hydrocarbons, and H₂, a quantity of the aromatics in the secondprocessed stream being less than that in the first processed stream; (c)removing, via a second reactor downstream of the aromatics separationdevice, at least some H₂ from the second processed stream, therebyproviding a recycle stream comprising CH₄ and C₂ hydrocarbons, aquantity of the H₂ in the recycle stream being less than that in thesecond processed stream; and (d) providing the recycle stream as atleast part of the input gas flow stream to the first reactor.
 12. Themethod of claim 11, wherein a composition of the first output stream isat least 50% aromatics.
 13. The method of claim 11, further comprising:repeating (a)-(d) at least two additional times, wherein after therepeating, at least 40% of an initial quantity of CH₄ is converted. 14.The method of claim 11, wherein: the converting of (a) is performed at atemperature greater than a temperature at which the removing of (c) isperformed.
 15. The method of claim 11, wherein: the first reactor has afirst gas flow volume, a second gas flow volume, and a first membraneseparating the first gas flow volume from the second gas flow volume,and the converting of (a) comprises flowing a first sweep gas throughthe second gas flow volume as the input gas flow stream is flowedthrough the first gas flow volume, the first sweep gas comprising O₂ oran oxygen-containing compound, such that hydrogen ions permeate throughthe first membrane from the first gas flow volume into the second gasflow volume and such that oxygen ions permeate through the firstmembrane from the second gas flow volume to the first gas flow volume;or the second reactor has a third gas flow volume, a fourth gas flowvolume, and a second membrane separating the third gas flow volume fromthe fourth gas flow volume, and the removing of (c) comprises flowing asecond sweep gas through the fourth gas flow volume as the secondprocessed stream is flowed through the third gas flow volume, the secondsweep gas comprising O₂ or an oxygen-containing compound, such that theat least some H₂ is removed by hydrogen ions permeating through thesecond membrane from the third gas flow volume into the fourth gas flowvolume and such that oxygen ions permeate through the second membranefrom the fourth gas flow volume into the third gas flow volume; or bothof the above.
 16. The method of claim 15, wherein: an exothermicreaction between the permeated hydrogen in the second gas flow volumeand the O₂ or the oxygen-containing compound in the second gas flowvolume heats the first reactor, such that the converting of (a) is anautothermal operation; or an exothermic reaction between the permeatedhydrogen in the fourth gas flow volume and the O₂ or theoxygen-containing compound in the fourth gas flow volume heats thesecond reactor such that the removing of (c) is an autothermaloperation; or both of the above.
 17. The method of claim 11, wherein:the removing of (b) comprises condensing the at least some aromatics;and the method further comprises storing the condensed aromatics for useor transport.
 18. The method of claim 11, wherein the first reactor orboth the first and second reactors comprise a respective catalyst. 19.The method of claim 18, wherein the catalyst of the first reactorcomprises Fe@SiO₂.
 20. The method of claim 11, wherein: the firstreactor or the second reactor comprises a membrane separating first andsecond gas flow volumes; the membrane comprises a perovskite-type oxidehaving a formula of M′Ce_(1-x-y)Zr_(x)M″_(y)O_(3-δ), where: M′ is aleast one of Sr and Ba; M″ is at least one of Ti, V, Cr, Mn, Fe, Co Ni,Cu, Nb, Mo, W, Pr, Nd, Pm, Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm and Yb; x isbetween 0.01 and 0.2, inclusive; and y is between 0.01 and 0.3,inclusive; and the membrane is provided on a porous support comprising aperovskite-type oxide having a formula of M′Ce_(1-z)Zr_(z)O_(3-δ), whereM′ is Sr or Ba, and z is between 0.01 and 0.3, inclusive.